DE102005013039A1 - Preparation of acrolein and/or acrylic acid involves feeding to first reaction zone at least two gaseous, propane-containing feed streams comprising fresh propane, and subjecting their propane fed to catalyzed dehydrogenation - Google Patents

Abstract

Acrolein and/or acrylic acid is prepared by feeding to a first reaction zone A at least two gaseous, propane-containing feed streams comprising fresh propane, and in reaction zone A, subjecting their propane fed to a heterogeneously catalyzed dehydrogenation to obtain a product gas mixture A comprising propane and propylene. Preparation of acrolein and/or acrylic acid involves feeding to a first reaction zone A at least two gaseous, propane-containing feed streams having fresh propane and in reaction zone A, subjecting their propane fed in this way to a heterogeneously catalyzed dehydrogenation to obtains a product gas mixture A comprising propane and propylene; conducting product gas mixture A out of reaction zone A and in a first separation zone A, removing a portion of the constituents other than propane and propylene in product gas mixture A and using remaining product gas mixture A' comprising propane and propylene; in a second reaction zone B to charge an oxidation reactor and, in the oxidation reaction, subjecting the propylene present in product gas mixture A' to a catalyst gas phase partial oxidation with molecular oxygen to give a product gas mixture B having acrolein or acrylic acid as the target product and also excess molecular oxygen; and conducting product gas mixture b out of reaction zone B and in a second separation zone, removing target product present in product gas mixture B, and of the remaining residual gas having unconverted propane, molecular oxygen and any unconverted propylene as one of the two-propane-containing feed streams into reaction zone A, where this recycling into he reaction zone A along the reaction path of the catalyzed dehydrogenation of propane in reaction zone A is effected such that at the feed point, >=5 mol.% of the propane fed to reaction zone A via the other feed streams has already been converted under dehydrogenating conditions in reaction zone A.

Description

• A) einer ersten Reaktionszone A wenigstens zwei gasförmige, Propan enthaltende Zufuhrströme zuführt, von denen wenigstens einer Frischpropan enthält, und in der Reaktionszone A ihr so zugeführtes Propan einer heterogen katalysierten Dehydrierung unterwirft, wobei ein Propan und Propylen enthaltendes Produktgasgemisch A erhalten wird,
• B) aus der Reaktionszone A Produktgasgemisch A herausführt und in einer ersten Trennzone A von den im Produktgasgemisch A enthaltenen, von Propan und Propylen verschiedenen, Bestandteilen wenigstens eine Teilmenge abtrennt, und dabei verbleibendes, Propan und Propylen enthaltendes, Produktgasgemisch A'
• C) in einer zweiten Reaktionszone B zur Beschickung wenigstens eines Oxidationsreaktors verwendet und in dem wenigstens einen Oxidationsreaktor das im Produktgasgemisch A' enthaltene Propylen einer (selektiven) heterogen katalysierten Gasphasenpartialoxidation mit molekularem Sauerstoff zu einem Acrolein, oder Acrylsäure, oder deren Gemisch als Zielprodukt sowie überschüssigen molekularen Sauerstoff enthaltenden Produktgasgemisch B unterwirft,
• D) aus der Reaktionszone B Produktgasgemisch B herausführt und in einer zweiten Trennzone B im Produktgasgemisch B enthaltenes Zielprodukt abtrennt und von dem dabei verbleibenden, nicht umgesetztes Propan, molekularen Sauerstoff sowie gegebenenfalls nicht umgesetztes Propylen enthaltenden Restgas wenigstens eine nicht umgesetztes Propan, molekularen Sauerstoff und gegebenenfalls nicht umgesetztes Propylen enthaltende Teilmenge als einen der wenigstens zwei Propan enthaltenden Zufuhrströme in die Reaktionszone A rückführt.

The present invention relates to a process for the preparation of acrolein, or acrylic acid or a mixture thereof of propane, wherein

• A) a first reaction zone A feeds at least two gaseous, propane-containing feed streams, of which at least one fresh propane contains, and subjects in the reaction zone A their supplied propane of a heterogeneously catalyzed dehydrogenation, wherein a propane and propylene-containing product gas mixture A is obtained
• B) leads out of the reaction zone A product gas mixture A and separates in a first separation zone A of the components contained in the product gas mixture A, propane and propylene different constituents, at least a subset, while remaining, propane and propylene containing, product gas mixture A '
• C) used in a second reaction zone B for feeding at least one oxidation reactor and in the at least one oxidation reactor contained in the product gas mixture A 'propylene a (selective) heterogeneously catalyzed gas phase partial oxidation with molecular oxygen to an acrolein, or acrylic acid, or their mixture as the target product and excess submitting molecular oxygen-containing product gas mixture B,
• D) removes product gas mixture B from reaction zone B and separates off the target product contained in product mixture B in a second separation zone B and at least one unconverted propane, molecular oxygen and, if appropriate, remaining unreacted propane, molecular oxygen and optionally unreacted propylene unreacted propylene-containing portion as one of the at least two propane-containing feed streams in the reaction zone A is recycled.

Acrylic acid is a major basic chemical that is used as a monomer to Production of polymers is used, for example in disperse distribution in aqueous Medium be used as a binder. Acrolein is a significant intermediate, for example, for the production of glutardialdehyde, methionine, folic acid and acrylic acid.

The in the preamble this method described for the production of acrolein, or acrylic acid or their mixture of propane is known (eg from WO 01/96270, of US 2003/0187299 A1 and from the documents DE-A 10245585 and DE-A 10246119 and the prior art cited in these documents).

The method for the separation of product contained in the product mixture B target product are characterized in that the target product z. B. converted by absorptive and / or condensing measures from the gaseous to the condensed phase. As an absorption agent come z. As water, aqueous solution and / or organic solvents into consideration. As part of this "condensation" of the target product normally remains a not in the condensed phase transient residual gas, which comprises the relatively difficult to be condensed constituents of the product gas mixture B. These are usually especially those components whose boiling point at atmospheric pressure (1 bar) ≤ -30 ° C (their total content of residual gas is generally ≥ 70% by volume, often ≥ 80% by volume and often ≥ 90% by volume), which includes primarily unreacted propane and excess in reaction zone B. In addition, the residual gas is usually inert diluent gases such as N ₂ , CO ₂ , noble gases (He, Ne, Ar, etc.), CO and to a lesser extent also acrylic acid, acrolein and / or H ₂ O (the water vapor content of the residual gas may be up to 25 vol .-%, often up to 20 vol .-% or up to 10 vol .-%, but often below 10 vol .-% or below 5 vol. % This residual gas forms (based on the amount of propane contained therein), the main amount (usually at least 80%, or at least 90%, or at least 95% or more) of the residual gas formed in the separation zone B and is therefore in This document also known as the main residual gas.

In particular, when the condensation of the target product is carried out by absorption by means of an organic solvent, usually at least one second unreacted propane and optionally unreacted propylene containing residual gas in the separation zone B (based on propane contained therein is its amount in comparison to the main residual gas amount normally much lower). This is due to the fact that the forming condensed phase to some extent unreacted propane and optionally unreacted propylene absorbs. In the further course of the extractive, distillative, crystallizing and / or desorptive separation of the target product from the condensed th phase, this unreacted propane and optionally propylene is usually recovered as part of at least one further gas phase and preferably recycled to the reaction zone A. This can be z. B. in a mixture with the main residual gas (in this document then called total residual gas). However, it can also take place in the form of independent gas streams to be recycled into the reaction zone A. The latter may be oxygen-free, or else oxygen-containing (secondary residual gas) (for example if it is obtained by stripping by means of air or at the top of a rectification column which has been flushed by means of air as a polymerization inhibitor).

Either Main residual gas, total residual gas and secondary residual gas form in the sense this invention in the reaction zone A traceable unreacted propane, molecular oxygen and optionally unreacted propylene containing residual gas. At molecular oxygen free in the separation zone B unreacted unreacted propane and optionally not Reacted propylene-containing residual gas can according to the invention in admixture with main residual gas and / or minor residual gas (i.e., as an ingredient, for example of total residual gas) and / or also independently (in this case acts it is not to the reaction zone A in the context of the invention recycled residual gas) be recycled to the reaction zone A. In the latter case, this recycling can be done without any restriction, i.e., e.g. B. as part of the starting reaction gas mixture the reaction zone A take place. Preferably, in the method according to the invention the total amount of unreacted in the separation zone B. Propane and optionally unreacted propylene containing gas flows recycled to the reaction zone A. subsets can (as will be explained in more detail later in the application) but if necessary also for other purposes, such as B. for energy production and / or syngas production and / or as a diluent gas be used in the reaction zone B.

In the abovementioned publications of the prior art, the return of unconverted propane, molecular oxygen and optionally unreacted propylene-containing residual gas to the reaction zone A in the same place as the supply of the other propane-containing feed streams into the reaction zone A should take place (ie as a constituent of the reaction gas starting mixture, cf., for example, 6 in US-2003/0187299 A1). A disadvantage of this procedure is that the molecular oxygen contained in the residual gas reduces the selectivity of the propene formation in the reaction zone A and increases the selectivity of by-product formation on the carbon oxides CO and CO ₂ as a result of partial full combustion of propane and / or propylene. This applies in particular if the residual gas to be recycled is main residual gas.

The DE-A 10211275 attempts to remedy the aforementioned problem by that you proceed as described above, but at the same time the product gas mixture A formed in the reaction zone A in two Aliquots of identical composition splits and one of the two Aliquots as a source of hydrogen in the reaction zone A returns. adversely However, that is the deactivation rate of the dehydrogenation catalysts not yet fully satisfied with this procedure can.

The The object of the present invention was therefore to be compared to the the above method improved process for the preparation of Acrolein, or acrylic acid or to provide their mixture of propane.

• A) einer ersten Reaktionszone A wenigstens zwei gasförmige, Propan enthaltende Zufuhrströme zuführt von denen wenigstens einer Frischpropan enthält, und in der Reaktionszone A ihr so zugeführtes Propan einer heterogen katalysierten Dehydrierung unterwirft, wobei ein Propan und Propylen enthaltendes Produktgasgemisch A (in der Regel wird das Produktgasgemisch A auch molekularen Wasserstoff enthalten) erhalten wird,
• B) aus der Reaktionszone A Produktgasgemisch A herausführt und in einer ersten Trennzone A von den im Produktgasgemisch A enthaltenen, von Propan und Propylen verschiedenen, Bestandteilen wenigstens eine Teilmenge abtrennt, und dabei verbleibendes, Propan und Propylen enthaltendes, Produktgasgemisch A'
• C) in einer zweiten Reaktionszone B zur Beschickung wenigstens eines Oxidationsreaktors verwendet und in dem wenigstens einen Oxidationsreaktor das im Produktgasgemisch A' enthaltene Propylen einer (selektiven) heterogen katalysierten Gasphasenpartialoxidation mit molekularem Sauerstoff zu einem Acrolein, oder Acrylsäure, oder deren Gemisch als Zielprodukt sowie überschüssigen molekularen Sauerstoff enthaltenden Produktgasgemisch B unterwirft,
• D) aus der Reaktionszone B Produktgasgemisch B herausführt und in einer zweiten Trennzone B im Produktgasgemisch B enthaltenes Zielprodukt abtrennt und von dem dabei verbleibenden, nicht umgesetztes Propan, molekularen Sauerstoff sowie gegebenenfalls nicht umgesetztes Propylen enthaltenden (Haupt- und/oder Neben- bzw. Gesamt-)Restgas wenigstens eine nicht umgesetztes Propan, molekularen Sauerstoff und gegebenenfalls nicht umgesetztes Propylen enthaltende Teilmenge (vorzugsweise wenigstens 50 Vol.-%, oder wenigstens 75 Vol.-% und ganz besonders bevorzugt die Gesamtmenge (jeweils individuell bezüglich des Gesamtrestgases, des Hauptrestgases und/oder des Nebenrestgases)) als einen der wenigstens zwei Propan enthaltenden Zufuhrströme in die Reaktionszone A rückführt, gefunden, das dadurch gekennzeichnet ist, dass diese Rückführung in die Reaktionszone A längs des Reaktionspfads der heterogen katalysierten Dehydrierung des Propans in der Reaktionszone A so erfolgt, dass an der Zufuhrstelle bereits wenigstens 5 mol-% des der Reaktionszone A über die anderen Zufuhrströme (davor) (insgesamt) zugeführten Propan in der Reaktionszone A dehydrierend umgesetzt sind (bezogen auf einmaligen Durchgang durch die Reaktionszone A).

Accordingly, a process for producing acrolein, or acrylic acid or its mixture of propane, wherein

• A) a first reaction zone A feeds at least two gaseous, propane-containing feed streams of which contains at least one fresh propane, and subjects in the reaction zone A their supplied propane of a heterogeneously catalyzed dehydrogenation, wherein a propane and propylene-containing product gas mixture A (usually the Product gas mixture A also contain molecular hydrogen),
• B) leads out of the reaction zone A product gas mixture A and separates in a first separation zone A of the components contained in the product gas mixture A, propane and propylene different constituents, at least a subset, while remaining, propane and propylene containing, product gas mixture A '
• C) used in a second reaction zone B for feeding at least one oxidation reactor and in the at least one oxidation reactor contained in the product gas mixture A 'propylene a (selective) heterogeneously catalyzed gas phase partial oxidation with molecular oxygen to an acrolein, or acrylic acid, or their mixture as the target product and excess molecular oxygen ent subjecting product gas mixture B to
• D) from the reaction zone B product gas mixture B and separated in a second separation zone B in the product gas mixture B target product and separated from the remaining, unreacted propane, molecular oxygen and optionally unreacted propylene containing (main and / or secondary or total -) residual gas at least one unreacted propane, molecular oxygen and optionally unreacted propylene-containing subset (preferably at least 50 vol .-%, or at least 75 vol .-% and most preferably the total amount (each individually with respect to the total residual gas, the main residual gas and / or the minor residual gas)) as one of the at least two propane-containing feed streams to the reaction zone A found, which is characterized in that this return to the reaction zone A along the reaction path of the heterogeneously catalyzed dehydrogenation of the propane in the reaction zone A is so, that at the feeder R at least 5 mol% of the reaction zone A to the other feed streams (before) (in total) supplied propane in the reaction zone A dehydrating reacted (based on a single pass through the reaction zone A).

Becomes in the method according to the invention not the total amount of total in the separation zone B (total) incurred (Main and / or secondary or total) residual gas in the reaction zone A returned, can this other subset as already mentioned z. For the purpose of energy or synthesis gas production or as a diluent gas in the reaction zone B continue to be used. In general, however, you will at least the half or two-thirds (i.e., 50 vol% or 66.6 vol%), preferably at least three quarters and most preferably the total amount of the aforementioned in the separation zone B accumulating (each individually with respect to the Main and / or secondary or Total) residual gas according to the invention in the Recycle reaction zone A. Falls in the Separation zone B only one unreacted propane, molecular oxygen and unreacted propylene-containing residual gas stream (this is common the rule), this is preferred according to the invention completely (if necessary less a guided into the reaction zone B as a diluent gas subset of identical composition) is recycled to the reaction zone A. He but can also be divided into two subsets of identical composition, and, as described above, only a partial amount in the reaction zone A returned and the other subset will be reused elsewhere. Falls in the separation zone B more than such a residual gas stream, so these residual gas streams (such already mentioned) together according to the invention (eg merged) or only occasionally or individually recycled to the reaction zone A. The percent feedback information refer to the residual gas in this document, in particular its Total amount (i.e., the sum of all residual gas flows). Usually exists (in particular main) residual gas (as already mentioned) to ≥ 70 vol .-%, often ≥ 80 vol .-% and often to 90 mol%, usually ≥ 95 mol%, or to ≥ 98 mol% of constituents whose boiling point at atmospheric pressure (1 bar) is ≤ -30 ° C. According to the invention, the Repatriation of in the separation zone B accumulating, molecular oxygen-containing, (Main and / or secondary or total) residual gas in the reaction zone A longitudinal the reaction path of catalytic dehydrogenation, of course, not only one, but also several in a row arranged supply points distributed.

Under Fresh propane is understood in this document propane, which is the reaction zone A has not gone through yet. As a rule, it will be crude propane (the preferably the specifications according to DE-A 10246119 and DE-A 10245585 Fulfills) be, in small quantities also of propane different components contains.

Under the reaction path in the reaction zone A is in this document the flowpath of the reaction zone A over the propane supplied from residual gas from separation zone B (other) feed streams through the reaction zone A depending from the dehydrogenating turnover (the conversion in the heterogeneously catalyzed Dehydration) of this propane.

The the reaction zone A supplied Reaction gas starting mixture or feed gas mixture should in this writing the sum of all with the fresh propane on the same Height of Reaction path in the reaction zone A of the same gases supplied be.

According to the invention, the recirculation of the (main and / or secondary or total) residual gas along the reaction path in the reaction zone A into the reaction zone A preferably takes place such that at least 10 or at least 15 mol%, preferably already at least, already exist at the feed point 20 or at least 25 mol%, and most preferably at least 30 or at least 35 mol% and most preferably at least 35 or at least 40 mol% of the propane in the reaction zone A fed to the reaction zone A via the other feed streams (previously) (in total) (dehydrogenating) reacted (sales Z _U ) are. In general, at the feed point of the return of the (main and / or secondary or total) residual gas in the Reak tion zone A in the process according to the invention less than 70 mol%, often less than 60 mol% and often ≤ 50 mol% of the reaction zone A via the other feed streams supplied propane in the reaction zone A already dehydrogenating reaction (conversion Z _U ).

In the process according to the invention, as recommended in DE-A 10211275, the product gas mixture A formed in the reaction zone A is divided into two subsets of identical composition and one of the two subsets is returned to the reaction zone A (especially if this recycling is part of the reaction zone A fed starting reaction gas mixture is carried out), then, instead of all the aforementioned turnover numbers Z _U , the turnover numbers Z _U ^KR defined as follows apply:

KGV is the ratio of recycled to the reaction zone A subset the product gas mixture A to in the reaction zone A formed total amount on product gas mixture A.

Of the Content of recycled to the reaction zone A, incurred in the separation zone B, (Main and / or secondary or total) residual gas of molecular oxygen becomes in the procedure according to invention normally 0.5% to 10% by volume, often 1 to 8% by volume and many times 2 to 5 vol .-% amount. He usually stirs in particular, therefore, that in the reaction zone B, an excess of molecular Oxygen (based on the stoichiometry the desired Target reaction) usually beneficial to the life of the oxidation catalysts and on the kinetics of the selective heterogeneous catalyzed therein Gas phase partial oxidation of propylene to acrolein, or to acrylic acid, or their mixture affects. In contrast to the circumstances in the reaction zone according to the invention A, the thermodynamic conditions in the reaction zone B is not substantially due to the molar reactant ratio influenced because the selective heterogeneously catalyzed gas phase partial oxidation of propylene to acrolein, or to acrylic acid, or to their mixture kinetic control subject.

The relationship the amount of propane that the reaction zone A via that from the separation zone B originated recycled (main and / or secondary or total) residual gas is fed to the amount of propane, that of the reaction zone A over total propane containing feed streams supplied is, is in the method according to the invention usually 0.1 to 10, or 0.5 to 5, preferably 0.5 to 1.5 or 3 to 5 amount.

In an embodiment will be the process of the invention in the reaction zone A next to that coming from the separation zone B. repatriated (main and / or secondary or total) residual gas contained propane as only second feed stream only one stream of fresh raw propane (contains next to Propane still small amounts of impurities). According to the invention preferred is this raw propane (as generally in this document) in the Documents DE-A 10245585 and DE-A 10246119 recommended specifications exhibit. The same applies the specification of the feed gas mixture of the reaction zone B in the process according to the invention. Of course can in the method according to the invention but also propane (and optionally propylene) containing exhaust gas streams from others as the method of the invention as feed streams be used in the reaction zone A (see WO 02/00587; Gas stream from section c) of this WO can also be the reaction zone B supplied become).

It is only essential to the invention that at the point (at the level) of the reaction zone A at which the feed of the (main and / or secondary or total) residual gas from the separation zone B into the reaction zone A takes place, at least 5 mol -% of the reaction zone A over all other feed streams supplied propane have been implemented within the to be carried out in the reaction zone A catalytic dehydrogenation. If the above condition is not met, the molar ratio of molecular hydrogen to molecular oxygen at the feed point is reduced, which promotes either the full combustion of the C ₃ hydrocarbons involved and or the service life of the dehydrogenation catalyst used reduces and / or necessarily the supply of molecular hydrogen an external source in the reaction zone A is required. This is presumably because both the C ₃ hydrocarbons involved and the dehydrogenation catalyst are better protected against full combustion and thermal (deactivating) decomposition on the dehydrogenation catalyst even in an atmosphere having an increased molar ratio of molecular hydrogen to molecular oxygen.

In contrast to the exothermic heterogeneously catalyzed oxydehydrogenation, which is enforced by oxygen present and in the intermediate no free hydrogen is formed (the torn from the hydrocarbon to be dehydrogenated hydrogen is torn off immediately as water (H ₂ O)) or after is to be understood, the heterogeneously catalyzed dehydrogenation according to the invention is to be understood as a ("conventional") dehydrogenation, the heat of reaction is endothermic in contrast to the oxydehydrogenation (as a subsequent step may be included exothermic hydrogen combustion in the process of the invention) and formed at the at least intermediate free molecular hydrogen This usually requires different reaction conditions and different catalysts than for oxydehydrogenation.

That is, the process of the invention normally proceeds under H ₂ evolution and this in particular until the return of the residual gas from the separation zone B in the reaction zone A, which is why the reaction gas mixture at the associated feed, based on the molar amount of propane contained in it usually one has higher molar hydrogen content, as the reaction zone A supplied reaction gas starting mixture.

The latter meets in the process of the invention preferably also to the reaction gas mixture, which is at the feed point by combining from the recycled in this residual gas and before that return forms the feed point present reaction gas mixture.

Otherwise has it is for the inventive method as cheap proven when within the reaction zone A, the molar ratio of in the reaction gas mixture contained propylene in the reaction gas mixture contained molecular hydrogen, the value 10, preferably the Value 5, better the value 3 and better still does not exceed the value 2. Particularly preferably, the aforementioned ratio moves at Values of 0.5 and 1 to 2, respectively.

For the rest, can the principle of heterogeneously catalyzed dehydrogenation in the reaction zone A (and thus the reaction zone A itself) in the process according to the invention as in the documents WO 01/96270, DE-A 3313573, DE-A 10131297, DE-A 10245585, DE-A 10246119 and DE-A 10211275 described become. That is, based on the single passage of the reaction zone A supplied Charge gas mixture through the reaction zone A, the reaction zone A through targeted heat exchange with outside the reaction zone A guided (fluid, i.e., liquid or gaseous) Heat transfer isothermal be designed. It can also be adiabatically that is, substantially without such targeted heat exchange with outside the reaction zone A guided Heat carriers are running. In the latter case, the gross monetary value, based on the one-off Passage of the reaction zone A supplied Reaktionsgasaus gangsgemischs through the reaction zone A, by grasping in the above Writings recommended and to be described in the following both endothermic (negative), or autothermal (essentially zero) or exothermic (positive). In accordance with the invention only has to additionally the repatriation of from the separation zone B originating residual gas in the reaction zone A inside as required by the invention respectively. Likewise the catalysts recommended in the abovementioned publications in the process according to the invention be applied.

In typically needed the heterogeneously catalyzed partial dehydrogenation of propane too Propylene comparatively high reaction temperatures. The achievable Sales is usually due to the thermodynamic equilibrium limited. Typical reaction temperatures are 300 to 800 ° C and 400 to 700 ° C. Per molecule on propylene dehydrogenated propane thereby becomes a molecule of hydrogen generated.

High temperatures and removal of the reaction product H ₂ favor the equilibrium position in the sense of the target product.

Since the heterogeneously catalyzed dehydrogenation reaction proceeds under volume increase, the conversion can be increased by lowering the partial pressure of the products. This can be achieved in a simple manner, for example by dehydration under reduced pressure and / or by admixing substantially inert diluent gases, such as, for example, water vapor, which is normally an inert gas for the dehydrogenation reaction. A dilution with water vapor as a further advantage, as a rule, a reduced coking of the catalyst used, since the water vapor reacts with coke formed according to the principle of coal gasification. In addition, water vapor can be used as diluent gas in the subsequent reaction zone B. However, steam can also be separated in a simple manner partially or completely from the product gas mixture of the dehydrogenation (the product gas mixture A) (for example by condensation), which opens up the possibility of further use of the modified product gas mixture (of the product gas mixture A ') in the process Reaction zone B to increase the proportion of diluent N ₂ . Further for the heterogeneously catalyzed propane dehydrogenation suitable diluents are, for example, CO, methane, ethane, CO _2, nitrogen and noble gases such as He, Ne and Ar. All of the diluents mentioned can be used either individually or in the form of very different mixtures be. It is advantageous that the said diluents are also suitable diluents in reaction zone B as a rule. In general, preference is given in the process according to the invention to inert diluents (that is to say less than 5 mol%, preferably less than 3 mol% and even better less than 1 mol%) in the particular reaction zone.

in principle come for the heterogeneously catalyzed propane dehydrogenation all in the state of Technically known dehydrogenation catalysts into consideration. They leave roughly divide into two groups. Namely in those that are more oxidic Nature are (for example, chromium oxide and / or alumina) and in those that consist of at least one on one, usually oxidic, carrier isolated, usually comparatively noble, metal (for Example platinum) exist. Among other things, so all dehydrogenation catalysts used in WO 01/96270, EP-A 731077, the DE-A 10211275, DE-A 10131297, WO 99/46039, US-A 4,788 371, EP-A-0 705 136, WO 99/29420, US-A 4 220 091, US-A 5 430 220, US-A 5,877,369, EP-A-0 117 146, DE-A 199 37 196, the DE-A 199 37 105 and DE-A 199 37 107 are recommended. In particular can both the catalyst according to Example 1, Example 2, Example 3, and Example 4 of DE-A 199 37 107 used become.

there are dehydrogenation catalysts containing 10 to 99.9% by weight Zirconia, 0 to 60 weight percent alumina, silica and / or Titanium dioxide and 0.1 to 10 wt .-% of at least one element of first or second main group, an element of the third subgroup, an element of the eighth subgroup of the Periodic Table of the Elements, Lanthanum and / or tin contain, with the proviso that the sum of the weight percentages 100 wt .-% results.

Especially also suitable in the examples and comparative examples Dehydrogenation catalyst used in this document.

As a general rule For example, the dehydrogenation catalysts may be catalyst strands (diameter typically 1 to 10 mm, preferably 1.5 to 5 mm; Length typically 1 to 20 mm, preferably 3 to 10 mm), tablets (preferably the same dimensions as in the strands) and / or catalyst rings (outer diameter and length each typically 2 to 30 mm or to 10 mm, wall thickness expediently 1 to 10 mm, or to 5 mm, or to 3 mm).

In As a rule, the dehydrogenation catalysts (especially those in this Example font used and those in DE-A 19937107 recommended (in particular the exemplary catalysts of this DE-A) so that they can both the dehydration of propane and catalyze the combustion of molecular hydrogen assets. The Hydrogen combustion proceeds in comparison with the dehydration of propane in the case of a competitive situation at the catalysts much faster.

to execution the heterogeneously catalyzed propane dehydrogenation come in principle all known in the art reactor types and process variants into consideration. Descriptions of such process variants included for example, all of them the dehydrogenation catalysts quoted prior art writings and the prior art cited at the beginning of this specification.

A comparatively comprehensive description of dehydrogenation processes suitable according to the invention 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.

Characteristic for the partial heterogeneously catalyzed dehydrogenation of propane is how already said that it is endothermic. That is, the for the Adjustment of the required reaction temperature as well as for the reaction needed Thermal energy) must either the reaction gas starting mixture in advance and / or in the course of the heterogeneous catalyzed dehydration supplied become. Optionally, the reaction gas mixture must the needed heat of reaction to withdraw oneself.

Further is it for heterogeneously catalyzed dehydrogenation of propane due to the high required reaction temperatures typical that in small quantities high-boiling high molecular weight organic Compounds, down to the carbon, are formed on the catalyst surface Separate and thereby deactivate selbige. To this detrimental This can be catalyzed by heterogeneously catalyzed side effects Dehydration at elevated Temperature over the catalyst surface to be passed, propane-containing reaction gas mixture with steam dilute become. Depositing carbon is among those given Partially or completely eliminated conditions according to the principle of coal gasification.

A different possibility, eliminating deposited carbon compounds is to the dehydrogenation catalyst from time to time at elevated temperature with an oxygen-containing gas (useful in the absence of hydrocarbons) to flow through and burn off the deposited carbon, so to speak. A extensive suppression The formation of carbon deposits is also possible because of the heterogeneously catalyzed propane to be dehydrogenated before it at elevated Temperature over led the dehydrogenation catalyst is added, molecular hydrogen.

Of course, there is also the possibility the heterogeneously catalysed to be dehydrogenated propane a mixture of Water vapor and molecular hydrogen to add. An addition of molecular hydrogen for heterogeneously catalyzed dehydrogenation of propane also reduces the unwanted formation of allenes (propadiene), Propyne and acetylene as by-products. Partial oxidation of such added Hydrogen can also provide the required heat of reaction.

A suitable reactor form for the heterogeneously catalyzed propane dehydrogenation in the reaction zone A. is the fixed bed tube or tube bundle reactor. That is, the Dehydrogenation catalyst is in one or in a bundle of Reaction tubes as a fixed bed. In the contact tube can according to the invention advantageous centered to run a second tube, the most diverse Heights (or only at one height) Has exit points (see WO 01/85333 and WO 01/85330), via the the residual gas from the separation zone B can be supplied. alternative can the in contact tube bottoms integrated contact tubes have interruption sections (spaces), in which the residual gas from the separation zone B can be added. The Reaction tubes are heated by the fact that in the reaction tubes surrounding space a gas, for example a hydrocarbon like Methane, burned. Cheap It is this direct form of contact tube heating only at first about 20 to 30% of the fixed bed apply and the remaining bed length by in the context of Combustion released radiant heat to the required reaction temperature warm. In this way is an approximate isothermal reaction regime reachable. Suitable reaction tube internal diameters are approximately 10 to 15 cm. A typical Dehydrierrohrbündelreaktor comprises 300 to 1000 reaction tubes. The temperature in the reaction tube interior moves in the range of 300 to 700 ° C, preferably in the range of 400 to 700 ° C. Advantageously, the reaction gas starting mixture is the Tube reactor supplied preheated to the reaction temperature. It is possible that the product gas mixture A, the reaction tube with a 50 to 100 ° C lower located Temperature leaves. These Output temperature can also be higher or at the same level lie. In the context of the aforementioned procedure is the use oxidic dehydrogenation catalysts based on chromium and / or alumina appropriate. Frequently becomes no diluent gas co-use, but of essentially sole raw propane go out as starting reaction gas. Also the dehydrogenation catalyst is usually undiluted applied.

Industrially one can in the reaction zone A several (eg., Three) such tube bundle reactors in parallel operate. It can if appropriate according to the invention one (or two) of these reactors are in dehydrogenation operation, while in a second (third) reactor, the catalyst feed is regenerated, without the operation in the reaction zone B suffers.

A such procedure is for example in the literature known BASF Linde propane dehydrogenation appropriate. It is according to the invention but important that the use of such a tube bundle reactor is sufficient.

Such a technique is also applicable to the so-called "steam active reforming (STAR) process" developed by Phillips Petroleum Co. (see, for example, US-A 4,902,849, US-A 4,996,387 and US-A 5) 389 342). As Dehydrierkataly capacitor in the STAR process with promoters containing platinum on zinc (magnesium) spinel is used as a carrier (see, for example, US-A 5,073,662). In contrast to the BASF-Linde propane dehydrogenation process, the propane to be dehydrogenated is diluted with steam in the STAR process. Typical is a molar ratio of water vapor to propane in the range of 4 to 6. The reactor outlet pressure is often 3 to 8 bar and the reaction temperature is suitably selected to 480 to 620 ° C. Typical catalyst (bed) loading with propane is 200 to 4000 h ^-1 (GHSV).

Under the load of a catalyst bed catalyzing a reaction step with reaction gas (starting) mixture, the amount of reaction gas (exit) mixture in standard liters (= Nl; the volume in liters that the corresponding reaction gas (outlet) mixture amount under normal conditions, ie, at 0 ° C and 1 atm), which is passed through one liter of catalyst bed per hour. The load may also be related to only one component of the reaction gas (outlet) mixture. Then it is the amount of this ingredient in Nl / lh that is passed through one liter of the catalyst bed per hour. Instead of Nl / lh, abbreviated "h ^-1 " is often used.

The heterogeneously catalyzed propane dehydrogenation can in the process according to the invention also be designed in a moving bed. For example, the catalyst moving bed be housed in a radial flow reactor. In the same moves the catalyst slowly from top to bottom, while the Reaction gas mixture flows radially. This procedure is for example in the so-called UOP Oleflex dehydrogenation applied. Since the reactors operated quasi adiabatically in this process be, it is expedient, several Operate reactors as a cascade in series (in typically up to four). Within the cascade, the residual gas fed from the separation zone B. become. This can lead to high differences in temperatures the reaction gas mixture at the reactor inlet and at the reactor outlet avoid (in the adiabatic mode of operation, the reaction gas starting mixture as a heat transfer medium, from its heat content the drop in the reaction temperature is dependent) and still appealing total sales achieve.

If the catalyst bed has left the moving bed reactor, it will fed to the regeneration and subsequently reused. As a dehydrogenation catalyst, for example, for this process a spherical one Dehydrogenation catalyst can be used, consisting essentially of Platinum on spherical alumina support consists. In the UOP variant, the propane to be dehydrogenated becomes Hydrogen added, to avoid premature catalyst aging. The working pressure is typically 2 to 5 atm. The hydrogen to propane ratio (the molar) is suitably 0.1 to 1. The reaction temperatures are preferably 550 to 650 ° C and the Residence time of the catalyst in a reactor is about 2 to 10 h elected.

at The fixed bed method described, the catalyst geometry also spherical, but also cylindrical (hollow or full) or otherwise geometric be designed.

When more possible Process variant for the heterogeneously catalyzed propane dehydrogenation in the process according to the invention describes Proceedings De Witt, Petrochem. Review, Houston, Texas, 1992 a, N1 the possibility a heterogeneously catalyzed propane dehydrogenation in a fluidized bed.

According to the invention can thereby z. B. two fluidized beds are operated side by side, of which one without negative impact on the overall process at times can be in the state of regeneration. As active material comes while chromium oxide on alumina used. The working pressure is typically 1 to 2 bar and the dehydrogenation temperature is usually at 550 to 600 ° C. The for the dehydration required heat is thereby introduced into the reaction system that the dehydrogenation catalyst on the reaction temperature is preheated. The above dehydrogenation is also known in the literature as the Snamprogetti-Yarsintez method.

alternative to the procedures described above, the heterogeneous catalyzed propane dehydrogenation in the process according to the invention also realized according to a process developed by ABB Lummus Crest (see Proceedings De Witt, Petrochem .. Review, Houston, Texas, 1992, P1).

at Both methods mentioned above, the residual gas addition according to the invention take place as a simple feed stream at the claimed reaction path path.

The previously described heterogeneously catalyzed dehydrogenation propane has in common that at propane conversions of ≥ 30 mol% (usually ≤ 60 mol%) operated (based on a single pass through the reaction zone and total supplied Propane). According to the invention of advantage however, that it is sufficient to have a propane conversion of ≥ 5 mol% to ≤ 30 mol% or ≤ 25 mol%, to achieve. This means, the heterogeneous invention catalyzed propane dehydrogenation can also occur in reaction zone A. at propane conversions from 10 to 20 mol% (the sales refer to a single pass through the reaction zone A). This stems among other things from the fact that the remaining amount of unreacted propane in the following Reaction zone B essentially acts as an inert diluent gas and according to the invention as a component of the resulting in the separation zone B residual gas largely lossless be recycled to the reaction zone A. can.

However, preference is given according to the invention to increased propane conversions in reaction zone A, since this reduces the formation of propionic acid by-products in reaction zone B. Such advantageous propane conversions are z. B. 30 or 40 to 50 or 60 mol% (based on a single pass through the reaction zone A and total propane supplied).

For the realization the aforementioned propane conversions is it cheap the heterogeneously catalyzed propane dehydrogenation at a working pressure from 0.3 to 3 bar. Furthermore, it is convenient the heterogeneously catalytically dehydrogenated propane with water vapor to dilute. So allows the heat capacity of the water on the one hand, part of the effect of the dehydrogenation endotherm and on the other hand reduces the dilution with Water vapor the Edukt- and product partial pressure, which is favorable the equilibrium position of the dehydration effect. It also affects the Wasserwampfmitverwendung, as already mentioned, advantageous to the life of noble metal-containing dehydrogenation catalysts. If necessary can as a further component also added molecular hydrogen become. The molar ratio from molecular hydrogen to propane is usually ≤ 5. The molar relationship from steam to propane can z. B. ≥ 0 to 30, suitably 0.1 to 2 and cheap 0.5 to 1. An advantage of a low-level approach Propane conversion is that with a single pass of the reaction gas only a comparatively low through the reaction zone A. heat is consumed and relatively low for sales Reaction temperatures are sufficient.

Also can the propane dehydrogenation, as already stated, according to the invention (quasi) Adiabatic and thereby be carried out endothermically. This is the Reaction gas starting mixture usually initially to a temperature of 500 to 700 ° C (or from 550 to 650 ° C) heated (for example, by direct firing of the surrounding wall). Adiabatic passage through the at least one catalyst bed then the reaction gas mixture is then depending on the conversion and dilution about 30 ° C Cool to 200 ° C. Behind the injection of the invention Residual gases from the separation zone B will be heated to some extent. A presence of water vapor as a heat carrier is also under the Viewpoint of adiabatic driving advantageously noticeable. lower Reaction temperatures allow longer service life the catalyst bed used. Higher reaction temperatures promote increased Sales.

in principle is the heterogeneously catalyzed propane dehydrogenation in the reaction zone A, independent whether adiabatic or isothermal, both in a fixed bed reactor as well as in a moving bed or fluidized bed reactor feasible.

Remarkably, can in the inventive method as reaction zone A also in adiabatic operation a single shaft furnace reactor be sufficient as a fixed bed reactor, the reaction gas mixture flows through axially and / or radially becomes.

in the the simplest case is a single closed one Reaction volume, for example, a container whose inner diameter 0.1 to 10 m, possibly also 0.5 to 5 m and in which the fixed catalyst bed on a carrier device (For example, a grid) is applied. That with catalyst charged reaction volume, which in adiabatic operation substantially thermally insulated is, is going to mean, Propane containing, reaction gas flows through axially. The catalyst geometry can be both spherical as well as annular or strand-shaped be. about in the catalyst bed introduced supply lines may originate from the separation zone B. Injected residual gas become. Since in this case the reaction volume by a very inexpensive Apparatus is to be realized, are all catalyst geometries, the have a particularly low pressure loss, to be preferred. These are mainly catalyst geometries leading to a large void volume to lead or structured, such as monoliths or Honeycombs. For realizing a radial flow of the propane-containing reaction gas For example, if the reactor is made of two sheaths, concentric nested, cylindrical gratings exist and the catalyst bed be arranged in the annular gap. In the adiabatic case, the metal shell would be optional again thermally isolated.

When Inventive catalyst charge for one heterogeneously catalyzed propane dehydrogenation are particularly suitable also those disclosed in DE-A 199 37 107, above all all by way of example disclosed catalysts.

After a prolonged period of use, the aforementioned catalysts can be regenerated in a simple manner, for example, by dilute air at first at initial temperatures of from 300 to 600 ° C., often at 400 to 550 ° C., first in regeneration stages with nitrogen and / or water vapor the catalyst bed passes. The catalyst (bed) load with regeneration gas (eg air) can be z. B. 50 to 10,000 h ^-1 and the oxygen content of the regeneration gas be 0.1 or 0.5 to 20 vol .-%.

In subsequent further regeneration stages, air can be used as regeneration gas under otherwise identical regeneration conditions. From an application point of view, it is advisable to purge the catalyst with inert gas (for example N ₂ ) before it is regenerated.

Then it is It is usually recommended, even with pure molecular hydrogen or with molecular inert by inert gas (vorzusgweise water vapor) Hydrogen (the hydrogen content should be ≥ 1 Vol .-%) in the otherwise same condition grid to regenerate.

The heterogeneously catalyzed propane dehydrogenation with comparatively low propane conversion (≦ 30 mol%) can in all cases be operated at the same catalyst (bed) loadings (both the reaction gas as a whole and the propane containing same in the same) as the variants with high propane conversion (> 30 mol%). This load of reaction gas may be, for example, 100 to 40,000 or 10000 h ^-1 , often 300 to 7000 h ^-1 , which is often about 500 to 4000 h ^-1 .

application point appropriate can be the reaction zone A in the process according to the invention as a tray reactor realize what the metered addition of the residual gas from the separation zone B allows between two hordes in a particularly simple manner.

This contains spatial successively more than one dehydrogenation catalyzing Catalyst bed. The catalyst bed number can be 1 to 20, suitably 2 to 8, but also 3 to 6 amount. Let with increasing Hordenzahl Increased increasingly easier propane conversions to reach. The catalyst beds are preferably radial or axial arranged one behind the other. Application technology is appropriate applied in such a tray reactor, the fixed catalyst bed type.

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

In expedient manner the reaction gas mixture is on its way from a catalyst bed for next Catalyst bed, for example by passing over heated with hot gases heat exchanger surfaces (z. B. ribs) or by passing through heated with hot fuel gases pipes, subjected in the Horden reactor of a Zwischenerhitzung.

Becomes the Horden reactor in the rest operated adiabatically, it is for Propane conversions ≤30 mol%, especially when using the described in DE-A 199 37 107 Catalysts, in particular the exemplary embodiments, sufficiently, the reaction gas mixture to a temperature of 450 up to 550 ° C preheated to pass into the dehydrogenation reactor and within the tray reactor to keep in this temperature range. That is, the entire propane dehydrogenation is so at extremely low Temperatures to realize what the service life of the fixed catalyst beds between two regenerations proves to be particularly favorable. For higher propane conversions will the reaction gas mixture expediently higher Temperatures preheated in the dehydrogenation reactor led (this can up to 700 ° C amount) and within the tray reactor in this elevated temperature range held.

It is even more clever to carry out the above described intermediate heating directly (autothermal mode). To this end, molecular oxygen is added to the reaction gas mixture either already before flowing through the first catalyst bed (then the reaction gas starting mixture should contain molecular hydrogen added) and / or between the subsequent catalyst beds. Thus, (as a rule catalyzed by the dehydrogenation catalysts themselves) a limited combustion of molecular hydrogen contained in the reaction gas mixture and formed in the course of the heterogeneously catalyzed propane dehydrogenation and / or added to the reaction gas mixture (it may also be expedient from an application point of view to insert catalyst beds in the tray reactor, which are charged with catalyst which specifically catalyzes (selectively) the combustion of hydrogen (as such catalysts are, for example, those of the documents US Pat. No. 4,788,371, US Pat. No. 4,886,928, US Pat. No. 5,430,209, US Pat. 5,530,171, 5,527,979 and 5,563,314, for example, such catalyst beds can be accommodated in an alternating manner with the beds containing dehydrogenation catalyst in the tray reactor.) The heat of reaction thus released thus enables quasi-autothermal (the Gross thermal tone is essentially zero) way a nearly isothermal Operation of heterogeneously catalyzed propane dehydrogenation. With increasing selected residence time of the reaction gas in the catalyst bed so a propane dehydrogenation is possible at decreasing or substantially constant temperature, which allows a particularly long service life between two regenerations.

In general, according to the invention, an oxygen feed as described above should be carried out so that the oxygen content of the reaction gas mixture, based on the amount of molecular hydrogen contained therein, is 0.5 to 50 or 30, preferably 10 to 25,% by volume. Both pure molecular oxygen or inert gas, for example CO, CO ₂ , N ₂ and / or noble gases, dilute oxygen, but in particular air, as well as nitrogen oxides come into consideration as oxygen source. The resulting combustion gases generally additionally dilute and thus promote heterogeneously catalyzed propane dehydrogenation.

in the Framework of the procedure according to invention forms the recycled from the separation zone B in the reaction zone A (main and / or secondary or total) residual gas a particularly rich Source of molecular oxygen, according to the invention the reaction gas mixture in the horde reactor usually at the earliest would be added after passing through the first catalyst bed.

The Isothermia of heterogeneously catalyzed propane dehydrogenation can be thereby further improve that one in the Horden reactor in the spaces between the catalyst beds closed, but before its filling conveniently but not necessarily evacuated, internals (for example, tubular), attaches. such Built-ins can be placed in the respective catalyst bed. These fittings Contain suitable solids or liquids above one evaporate or melt at a certain temperature while consuming heat and condense again where this temperature falls below and heat release.

Needless to say, let yourself the reaction zone A of the method according to the invention also as in DE-A 10211275 described (as "loop variant") realize the forms an integral part of this patent application.

• a) das Produktgasgemisch A in zwei Teilmengen identischer Zusammensetzung aufteilt und eine der beiden Teilmengen (als Dehydrierkreisgas) als einen der wenigstens drei Propan enthaltenden Zufuhrströme in die Reaktionszone A rückführt (vorzugsweise als Bestandteil des Beschickungsgasgemischs (Reaktionsgasausgangsgemischs) der Reaktionszone A) und von der anderen Teilmenge an Produktgasgemisch A (vorzugsweise die gesamte andere Teilmenge) in der ersten Trennzone A erfindungsgemäß weiterbehandelt, und
• b) die Zuführung des (Haupt- und/oder Neben- bzw. Gesamt-)Restgases aus der Trennzone B in die Reaktionszone A längs des Reaktionspfads der heterogen katalysierten Dehydrierung des Propans so vornimmt, dass an der Zufuhrstelle bereits wenigstens 5 mol-%, oder wenigstens 10 mol-%, vorzugsweise wenigstens 15 mol-%, oder wenigstens 20 mol-%, besonders bevorzugt wenigstens 25 mol-%, oder wenigstens 30 mol-%, ganz besonders bevorzugt wenigstens 35 mol-%, oder wenigstens 40 mol-%, und am Besten wenigstens 45 mol-%, oder wenigstens 50 mol-% (in der Regel jedoch weniger als 70 mol-%, häufig weniger als 60 mol-%, und oft ≤ 50 mol-%) des der Reaktionszone A über die anderen Zuführströme (insgesamt) zugeführten Propan in der Reaktionszone A bereits dehydrierend umgesetzt sind (besonders bevorzugte Verfahrensvarianten sind dabei wiederum jene, bei denen die vorgenannten Umsatzzahlen Z_U jeweils durch Z_U ^Kr = (Z_U/1 + KGV) ersetzt sind). Dabei kann die Zuführung des (Haupt- und/oder Neben- bzw. Gesamt-)Restgases aus der Trennzone B in die Reaktionszone A an einer Stelle, oder auch auf mehrere hintereinander angeordnete Stellen verteilt erfolgen. Zweckmäßig wird man wenigstens 10 Vol.-%, bzw. 20 Vol.-% des Produktgasgemischs A in die Reaktionszone A rückführen. Vorteilhaft wird die als Kreisgas in die Reaktionszone A rückgeführte Menge an Produktgasgemisch A jedoch nicht mehr als 90 Vol.-%, bzw. 80 Vol.-% des Produktgasgemischs A betragen. D.h., die als Kreisgas in die Reaktionszone A rückgeführte Teilmenge des Produktgasgemischs A kann somit z. B. 20 bis 80 Vol.-%, oder 30 bis 70 Vol.-%, oder 40 bis 60 Vol.-% oder auch 50 Vol.-% des Produktgasgemischs A betragen. Besonders vorteilhaft ist der Mengenanteil 50 bis 70 Vol.-%.

That is, an embodiment of the method according to the invention is a method in which the reaction zone A at least three gaseous, propane-containing feed streams, at least one of which contains fresh propane and at least one recycled from the separation zone B in the reaction zone A, molecular oxygen, not reacted propane and (in the reaction zone B) optionally unreacted propylene containing (main and / or secondary or total) residual gas, and in the reaction zone A their supplied propane of a heterogeneously catalyzed (partial) dehydrogenation subjects (come in principle all the heterogeneously catalyzed dehydrogenation processes mentioned and stated in this document being considered), a propane-containing propylene-containing product gas mixture A being obtained with the proviso that

• a) the product gas mixture A is divided into two subsets of identical composition and one of the two subsets (as Dehydrierkreisgas) as one of the at least three propane-containing feed streams in the reaction zone A recycles (preferably as part of the feed gas mixture (reaction gas starting mixture) of the reaction zone A) and of the other Subset of product gas mixture A (preferably the entire other subset) further treated in the first separation zone A according to the invention, and
• b) the supply of the (main and / or secondary or total) residual gas from the separation zone B in the reaction zone A along the reaction path of the heterogeneously catalyzed dehydrogenation of propane so makes that at the feed point already at least 5 mol%, or at least 10 mol%, preferably at least 15 mol%, or at least 20 mol%, particularly preferably at least 25 mol%, or at least 30 mol%, very particularly preferably at least 35 mol%, or at least 40 mol% %, and most preferably at least 45 mol%, or at least 50 mol% (but usually less than 70 mol%, often less than 60 mol%, and often ≤ 50 mol%) of the reaction zone A the other feed streams (total) supplied propane in the reaction zone A are already implemented dehydrating (particularly preferred process variants are again those in which the aforementioned turnover numbers Z _{U are} each replaced by Z _U ^Kr = (Z _U / 1 + KGV)). In this case, the supply of the (main and / or secondary or total) residual gas from the separation zone B in the reaction zone A at one point, or even distributed over several locations arranged in succession. It is expedient to recycle at least 10% by volume, or 20% by volume, of the product gas mixture A into the reaction zone A. However, the amount of product gas mixture A recirculated as recycle gas into reaction zone A is advantageously not more than 90% by volume, or 80% by volume, of product gas mixture A. That is, the recycled as cycle gas in the reaction zone A subset of the product gas mixture A can thus z. B. 20 to 80 vol .-%, or 30 to 70 vol .-%, or 40 to 60 vol .-% or 50 vol .-% of the product gas mixture A amount. Particularly advantageous is the proportion of 50 to 70 vol .-%.

One possibility for heating the reaction gas starting mixture supplied to reaction zone A to the reaction temperature required for heterogeneously catalyzed propane dehydrogenation in reaction zone A is also to add molecular hydrogen to it and to oxidize it by means of molecular oxygen, for example to suitable specific combustion catalysts (for example referred to in this document), to burn (for example, by simply passing and / or passing), and by means of the combustion heat thus released to cause the heating to the desired reaction temperature. The resulting combustion products, such as CO ₂ , H ₂ O, and the N ₂ accompanying the molecular oxygen required for combustion, advantageously form inert diluent gases.

Becomes along the Reaction pathway of heterogeneously catalyzed dehydrogenation in the Reaction zone A in advance of the recycling of molecular oxygen containing residual gas from the separation zone B another molecular Oxygen-containing gas is passed into the reaction zone A is, independently of the specific embodiment of the reaction zone A, according to the invention expediently respected that in such a way supplied in advance molecular oxygen, the hydrogen content in Reaktionsgasge mix in the reaction zone A to the feed of the recirculating molecular oxygen-containing residual gas from the separation zone B not lowered too much.

According to the invention that From the separation zone B in the reaction zone A recycled residual gas and the only the reaction zone A supplied be molecular oxygen-containing gas (regardless of the concrete embodiment of the reaction zone A). This is not the rule of the procedure according to the invention.

• I. dem ersten Abschnitt der Reaktionszone A wenigstens einen, gasförmiges Propan enthaltenden Zufuhrstrom zuführt, der Frischpropan enthält, und in diesem ersten Abschnitt der Reaktionszone A das ihm so zugeführte Propan einer heterogen katalysierten Dehydrierung unterwirft, wobei ein Propan, Propylen und molekularer Wasserstoff enthaltendes Produktgasgemisch A* erhalten wird, bei dessen Erhalt wenigstens 5 mol-%, oder wenigstens 10 mol-%, vorteilhaft wenigstens 15 mol-%, oder wenigstens 20 mol-%, besonders vorteilhaft wenigstens 25 mol-%, oder wenigstens 30 mol-%, ganz besonders vorteilhaft wenigstens 35 mol-%, oder wenigstens 40 mol-%, und noch besser wenigstens 45 mol-%, oder wenigstens 50 mol-% (in der Regel jedoch weniger als 70 mol-%, oder weniger als 60 mol-%, oder ≤ 50 mol-%) des dem ersten Abschnitt der Reaktionszone A (insgesamt) zugeführten Propan in diesem ersten Abschnitt (der Reaktionszone A) dehydrierend (heterogen katalysiert) umgesetzt worden sind und das vorzugsweise, bezogen auf die darin enthaltene molare Menge an Propan, eine größere molare Menge an molekularem Wasserstoff enthält, als das dem ersten Abschnitt der Reaktionszone A zugeführte Reaktionsgasausgangsgemisch;
• II. dem Produktgasgemisch A* anschließend molekularen Sauerstoff, nicht umgesetztes Propan und gegebenenfalls (in der Reaktionszone B) nicht umgesetztes Propylen enthaltendes (Haupt- und/oder Neben- bzw. Gesamt-)Restgas aus der Trennzone B zuführt (erfindungsgemäß vorteilhaft wenigstens die Hälfte, bevorzugt wenigstens zwei Drittel oder drei Viertel und ganz besonders bevorzugt die Gesamtmenge (gegebenenfalls abzüglich einer als Verdünnungsgas in die Reaktionszone B rückgeführten Teilmenge identischer Zusammensetzung) des vorgenannten in der Trennzone B anfallenden (jeweils individuell bezogen auf Haupt- und/oder Neben- bzw. Gesamt-)Restgases) und das dabei resultierende Reaktionsgasgemisch A* dem zweiten Abschnitt der Reaktionszone A zuführt, und
• III. in diesem zweiten Abschnitt der Reaktionszone A unter Ausbildung des Propan und Propylen enthaltenden Produktgasgemischs A im Reaktionsgasgemisch A* enthaltenen molekularen Sauerstoff mit im Reaktionsgasgemisch A* enthaltenem molekularen Wasserstoff heterogen katalysiert zu Wasser verbrennt (vorteilhaft wenigstens 25 mol-%, bevorzugt wenigstens 50 mol-%, besonders bevorzugt wenigstens 75 mol-%, besser wenigstens 90 mol-% und am Besten wenigstens 95 mol-% der im Reaktionsgasgemisch A* enthaltenen Menge an molekularem Sauerstoff) und im Reaktionsgasgemisch A* enthaltenes Propan gegebenenfalls heterogen katalysiert zu Propylen dehydriert (in der Regel weniger als 40 mol-% bzw. weniger als 30 mol-%, meist weniger als 20 mol-%, häufig weniger als 10 mol-%, und oft weniger als 5 mol-% des im Reaktionsgasgemisch A* enthaltenen Propan). Vorgenannte Mengenverhältnisse können durch Anpassen der Ausmaße des zugehörigen Katalysatorbetts eingeregelt werden;
• IV. dann Produktgasgemisch A aus der Reaktionszone A herausgeführt und wie beschrieben in der ersten Trennzone A des erfindungsgemäßen Verfahrens erfindungsgemäß weiterbehandelt (gegebenenfalls nicht in die erste Trennzone A geführtes Produktgasgemisch A kann z. B. zur Energieerzeugung verbrannt und/oder zur Erzeugung von Synthesegas etc. verwendet werden).

In accordance with the invention, the reaction zone A will be designed and operated with the proviso that it consists of a first and a second section (hereinafter referred to as "two-section variant"), wherein

• I. supplying to the first section of the reaction zone A at least one feed stream containing gaseous propane containing fresh propane and subjecting in said first section of reaction zone A the propane thus supplied to a heterogeneously catalyzed dehydrogenation, wherein a propane, propylene and molecular hydrogen containing product gas mixture A * is obtained, upon receipt of which at least 5 mol%, or at least 10 mol%, advantageously at least 15 mol%, or at least 20 mol%, more preferably at least 25 mol%, or at least 30 mol%, most preferably at least 35 mol%, or at least 40 mol%, and more preferably at least 45 mol%, or at least 50 mol% (but usually less than 70 mol%, or less than 60 mol%). , or ≤ 50 mol%) of the propane fed to the first section of the reaction zone A (in total) in this first section (the reaction zone A) has been reacted by dehydrogenation (heterogeneously catalyzed), and this is preferred Sage, based on the molar amount of propane contained therein, contains a larger molar amount of molecular hydrogen than the first portion of the reaction zone A supplied starting reaction gas mixture;
• II. The product gas mixture A * then molecular oxygen, unreacted propane and optionally (in the reaction zone B) unreacted propylene containing (main and / or secondary or total) residual gas from the separation zone B supplies (according to the invention advantageously at least half , preferably at least two-thirds or three-fourths and very particularly preferably the total amount (optionally less than a diluent in the reaction zone B recycled partial amount of identical composition) of the aforementioned occurring in the separation zone B (each individually based on main and / or secondary or Gesamt-) residual gas) and the resulting reaction gas mixture A * to the second section of the reaction zone A supplies, and
• III. in this second section of the reaction zone A with formation of the propane and propylene-containing product gas mixture A in the reaction gas mixture A * contained molecular oxygen heterogeneously catalyzed with water contained in the reaction gas mixture A * molecular hydrogen (advantageously at least 25 mol%, preferably at least 50 mol% , particularly preferably at least 75 mol%, more preferably at least 90 mol% and most preferably at least 95 mol% of the amount of molecular oxygen contained in the reaction gas mixture A *) and propane contained in the reaction gas mixture A * optionally dehydrogenated catalytically heterogeneously to give propylene (in US Pat Usually less than 40 mol% or less than 30 mol%, usually less than 20 mol%, often less than 10 mol%, and often less than 5 mol% of the propane contained in the reaction gas mixture A *). The aforementioned proportions can be adjusted by adjusting the dimensions of the associated catalyst bed;
• IV., Product gas mixture A is then led out of reaction zone A and further treated according to the invention in the first separation zone A of the process according to the invention (product gas mixture A which may not be introduced into first separation zone A may be burnt for energy production and / or synthesis gas, etc . be used).

When Catalysts are coming for the second section of the reaction zone A while all those catalysts into consideration, which catalyzes in this document also for the heterogeneous Dehydration be recommended and especially for the first Section of the reaction zone A are suitable because they, as input already mentioned in this document, usually also the combustion of molecular hydrogen too catalyze assets (This applies in particular to the catalysts of DE-A 19937107 (especially those exemplified) too; in a competitive situation between heterogeneously catalyzed propane dehydrogenation and heterogeneous catalyzed hydrogen combustion, the latter is usually essential faster and dominates the former). Of course, come for the second Section of the reaction zone A but also such catalysts in Consider that specific to tailoring the selective combustion of molecular hydrogen are. Such catalysts are z. B. those of the documents US-A 4788371, US-A 4886928, US-A 5430203, US-A 5530171, US-A 5527979 and US-A 5,563,314th

Basically the first portion of reaction zone A is both isothermal and be designed adiabatically. In the latter case, one is on the one time Passage of the first section of the reaction zone A added (Fed) Reaction gas starting mixture through the first section of the reaction zone A related endothermic to autothermal gross heat tinting over the first reaction section the reaction zone A is preferred.

The Reaction conditions in the second section of the reaction zone A (reaction temperature (eg 400 or 500 to 800 or 700 ° C.), reaction pressure (eg 1 to 10, or to 5, or to 3 bar) and load of the catalyst bed with reaction gas mixture (eg 500 (or less) to 80,000 (or more) Nl / l · h) can basically be similar to be selected in the first section of the reaction zone A.

in principle can also the second section of the reaction zone A isothermal or be designed adiabatically. In the latter case, one is on the one time Passage of the second section of the reaction zone A added (Fed) Reaction gas starting mixture (the reaction gas mixture A *) by the second section of the reaction zone A related exothermic Grosswärtenönung on the second reaction section of the reaction zone A preferred.

According to the invention preferred become both the first and the second section of the reaction zone A adiabatic designed, the combination "endothermic to autothermal Gross warmth "in the first reaction section and "exothermic Gross heat tint in the second reaction section is particularly preferred.

Especially advantageous is an embodiment of the first section of Reaction zone A in Horden structure, which is preferred according to the invention is operated in an adiabatic embodiment endothermic to autothermic. The horde structure contains one or more spatially successive catalyst beds equipped with dehydrogenation catalyst. The number of catalyst beds may be 1 to 20, conveniently 2 to 8, in particular 3 to 6 (for high Dehydrierumsätze in the first Section of the reaction zone A is a large number of catalyst beds advantageous).

The Catalyst beds are preferably radially or axially from the reaction gas flows through.

in the In the simplest case, the fixed catalyst beds are in a reactor axially or in the annular gaps of centrally nested cylindrical gratings arranged one behind the other. It can it is the reactor z. B. act to a shaft furnace. The implementation of the first section of the reaction zone A in a single shaft furnace is also possible as the realization of the entire reaction zone A in a tray reactor.

Would that be Reaction gas before and / or after entering the first section the reaction zone A at any point containing molecular oxygen Gas would be added the reaction gas in the structured in hordes first section of Reaction zone A on its way from one catalyst bed to the next catalyst bed appropriate one Subjected to intermediate heating, z. B. by passing over heated with hot gases heat exchanger surfaces, for. B. heat exchanger fins, or by passing through internals heated with hot fuel gases, z. B. pipes.

In the context of the method according to the invention, however, the above-described intermediate heating is preferably carried out at least partly directly. For this purpose, a gas containing molecular oxygen (preferably air or mixtures of oxygen and inert gases such as CO ₂ , N ₂ and noble gases, or pure O ₂ ) is added to the reaction gas either already before flowing through the first catalyst bed and / or between the subsequent catalyst beds to a limited extent ,

At one, or more, or optionally all, catalyst beds is thus added to the reaction gas in advance to a limited extent and / or formed during the dehydrogenation of molecular hydrogen burned. The released heat of reaction thus enables adiabatic design a substantially autothermal mode of operation of the first section the reaction zone A. Based on the reaction gas contained Amount of molecular hydrogen should be added to the molecular Oxygen amount 0.5 to 50 or 30 vol .-% (preferably 10 to 25% by volume). Due to the high heat of combustion of the Hydrogen is a partial combustion of comparatively small amounts same for the aforementioned purpose is usually sufficient.

In one embodiment of the invention, an intermediate feed of oxygen-containing gas is optionally carried out before each horde of the tray reactor. In a further embodiment of the method according to the invention, the feed of oxygen-containing gas takes place before each horde except the first horde. In a further embodiment of the process according to the invention, a bed of specific oxidation catalyst suitable for purposes of H ₂ oxidation is present behind each oxygen feed point, followed by a bed of dehydrogenation catalyst. If necessary, external molecular hydrogen (in pure form or diluted with inert gas) may additionally be added upstream of each horde. In a less preferred embodiment, the catalyst beds may also contain mixtures of dehydrogenation and H ₂ oxidation catalysts.

However, it is essential to the invention that the process of the invention does not require any feed of external molecular hydrogen in a necessary manner. The dehydrogenation temperature in the tray structure (in the tray reactor) is generally 400 to 800 ° C, the pressure generally 0.2 to 10 bar, preferably 0.5 to 4 bar and particularly preferably 1 to 3 bar. The total gas load (GHSV) is usually 500 to 10000 h ^-1 , in high load mode also up to 80000 h ^-1 , regularly at 30,000 to 40,000 h ^-1 .

presence from water vapor is beneficial for the Catalyst life.

partial pressure the reactants by reducing pressure, burning in the dehydration formed hydrogen and / or inert dilution, high temperatures and size Horde numbers promote the Sales of dehydrogenation within the first section of the reaction zone A.

For the second Section of the reaction zone A is preferably the same type of reactor as for the first section of the reaction zone A applied. However, he contains in usually just a catalyst bed. Preferably, in Both sections use simple shaft furnaces.

The molar ratio from molecular oxygen to molecular hydrogen can (in the reaction gas mixture A *) in the second section of the reaction zone A 1: 2 to 1:10, preferably 1: 2 to 1: 4 amount. That is, in the second section of the reaction zone A can essentially all of the reaction mixture contained in the A * molecular oxygen with molecular hydrogen burned to water so that the product gas mixture A formed thereby largely free of molecular oxygen. For example, it can do that formed product gas mixture A less than 5 vol .-%, or less as 3 vol.%, or less than 1 vol.%, often even less than Contain 0.5 or less than 0.2% by volume of molecular oxygen. Preference is given to a vanishing content of molecular oxygen.

Of the first section of the reaction zone A and the second section of the Reaction zone A can both in separate reactors, as well as in a single Reactor (eg in a tray reactor, eg in shaft furnace design) housed be.

In even more elegant according to the invention Way lets the reaction zone A as a combination of the above-described "loop variant" and the "two-section variant" designed as follows and operate.

there This will be done as described above in the "two-section variant" in the second section of Reaction zone A formed product gas mixture A additionally in split two aliquots of identical composition and one of the both subsets (as Dehydrierkreisgas) as another propane containing feed stream into the first section of the reaction zone A return (preferred as part of the reaction gas starting mixture of the first section the reaction zone A) and of the other subset of product gas mixture A (preferably the entire other subset) in the first separation zone Continue treatment according to the invention.

The Repatriation of Product gas mixture A in the first section of the reaction zone A. is according to the invention insofar advantageous than that formed in the second section of the reaction zone A. Product gas mixture A due to the combustion of molecular oxygen with molecular hydrogen on the one hand normally (at least for adiabatic exothermic operation of the second section of the reaction zone A) an increased Temperature and, on the other hand, also as a result of oxygen combustion, a low oxygen content and a significant water vapor content having. The hydrogen contents are usually low as well. Especially with smaller Dehydrierumsätzen in the first section of the Reaction zone A can the reaction gas mixture A * in advance of its supply in the second section of the reaction zone A supplementary external molecular Hydrogen (that is molecular hydrogen that is neither a constituent of recycled to the reaction zone A circulating gas is still in the reaction zone A (or in one of the other reaction / separation zones of the method according to the invention) self-formed) may be added, then in the second section the reaction zone A z. B. the desired temperature of the reaction gas mixture A (and elimination of molecular oxygen in selbigem) too to reach.

In accordance with the aforesaid is suitable in such a way as part of the reaction gas starting mixture in the first section the reaction zone A recycled product gas mixture A in a special way to that the first section of the reaction zone A supplied Reaction gas starting mixture to warm to reaction temperature, without at the same time the reaction gas starting mixture already contained molecular hydrogen thermodynamically or by contained molecular To charge oxygen kinetically. The product contained in the recirculated product gas mixture A. Promotes water vapor the heterogeneously catalyzed dehydrogenation in the first section of the Reaction zone A in addition and makes a separate supply of external water vapor usually unnecessary.

that is, the procedure of the invention impresses u.a. in that it is a mode of operation according to the invention with satisfactory service life, without it being in any of the reaction or separation zones of the feed from external water vapor (this is water vapor that does not in any of the Reaction / separation zones of the inventive method is formed) requirement.

Of the guided in the first separation zone A. Part of the product gas mixture formed in the second section of the reaction zone A. A is suitable because of its elevated temperature at the same time in an excellent way, by indirect heat exchange in a gas cooler both first that of the reaction zone A supplied Fresh propane and that from the separation zone B in the reaction zone A recycled residual gas (or just one of the two gases) (essentially to the desired reaction temperature) to warm, and at the same time in for the needs in the first separation zone A advantageously cool.

According to the invention is appropriate at least 10 vol.% or 20 vol.% of that in the second section formed product gas mixture A in the first section of the reaction zone A return. According to the invention advantageous is called the recycle gas in the first section of the reaction zone A returned subset of Product gas mixture A but not more than 90% by volume or 80% by volume be. That is, as a recycle gas in the first section of the reaction zone A returned subset of the product gas mixture A can in the aforementioned procedure z. 20 to 80 vol.%, Or 30 to 70 vol.%, Or 40 to 60 vol.%, or 50 vol .-% of the product gas mixture A formed. Particularly advantageous is a (dehydrogen) cycle gas fraction of 50 to 70% by volume.

In very elegant way the repatriation of the aforementioned product gas mixture A in the first section of the reaction zone A shape as follows.

First, the product gas mixture A *, the molecular oxygen, propane and optionally (in the reaction zone B) unreacted propylene containing (main and / or secondary or total) residual gas from the separation zone B according to the principle of one with this residual gas as Jet-driven jet pump ( 1 DE-A 10211275 explains this principle, which in this document should also include the principle of the ejector jet nozzles), wherein the conveying direction of the jet through a drive nozzle ( 1 ) via a mixing section ( 2 ) and a diffuser ( 3 ) relaxed propulsion jet in the second section of the reaction zone A and the suction (suction) of the suction nozzle ( 4 ) in the direction of the first section of the reaction zone A and the connection "suction port mixing section diffuser" forms the sole connection between the two sections of the reaction zone A.

In the suction nozzle, a negative pressure is thereby generated, sucked through from the first section of the reaction zone A product gas mixture A * and transported with simultaneous mixing with the propulsion jet as reaction gas mixture A * through the mixing section (the mixing tube) via the diffuser and in the two section of the reaction zone A (the numeric addresses refer to the 1 this writing).

In necessarily one will choose the pressure of the propulsion jet (the Pressure adjustment is usually by means of a centrifugal compressor or (in rarer cases) by means of a blower (usually a low pressure ratio axial compressor) Final pressure to outlet pressure)), that the pressure of the second section of the reaction zone A forming product gas mixture A above the pressure in the first section of the reaction zone A forming product gas mixture A * is. A simple pipe connection between the exit of the second section of the reaction zone A. and the entrance to the first section of the reaction zone A, the it is useful as Tee with flow divider is a subset of the product gas mixture A in the first separation zone A forwarding is then usually sufficient to get an in the first portion of the reaction zone A returning (the pressure gradient following) circular flow (a Kind of natural Circulation) of the other part of the product mixture A.

• I. dem ersten Abschnitt (0) der Reaktionszone A wenigstens einen, gasförmiges Propan enthaltenden Zufuhrstrom zuführt, der Frischpropan (5) enthält, und in diesem ersten Abschnitt der Reaktionszone A das ihm zugeführte Propan einer heterogen katalysierten (partiellen) Dehydrierung unterwirft, wobei ein Propan, Propylen und molekularen Wasserstoff enthaltendes Produktgasgemisch A* erhalten wird, bei dessen Erhalt wenigstens 5 mol-%, oder wenigstens 10 mol-%, vorteilhaft wenigstens 15 mol %, oder wenigstens 20 mol-%, besonders vorteilhaft wenigstens 25 mol-%, oder wenigstens 30 mol-%, ganz besonders vorteilhaft wenigstens 35 mol-%, oder wenigstens 40 mol-%, und noch besser wenigstens 45 mol-%, oder wenigstens 50 mol-% (in der Regel jedoch weniger als 70 mol-%, oder weniger als 60 mol-%, oder ≤ 50 mol-%) des dem ersten Abschnitt der Reaktionszone A (insgesamt) zugeführten Propan in diesem ersten Abschnitt dehydrierend umgesetzt worden sind (bei Schlaufenfahrweise in der Reaktionszone A sind bevorzugte Verfahrensvarianten dabei wiederum jene, bei denen die vorgenannten Umsatzzahlen Z_U jeweils durch Z_U ^Kr = (Z_U/1 + KGV) ersetzt sind);

dabei wird der erste Abschnitt der Reaktionszone A im wesentlichen adiabat betrieben (d.h. über außerhalb des ersten Abschnitts der Reaktionszone A befindliche Wärmeträger wird dem ersten Abschnitt der Reaktionszone A im wesentlichen weder (gezielt) Wärme zugeführt noch entnommen);
bezogen auf einmaligen Durchgang des Beschickungsgasgemischs des ersten Abschnitts der Reaktionszone A durch den ersten Abschnitt der Reaktionszone A ist die Bruttowärmetönung über den ersten Abschnitt der Reaktionszone A endotherm (negativ) bis autotherm (im wesentlichen Null);
vorzugsweise wird der erste Abschnitt der Reaktionszone A in Hordenstruktur ausgeführt;
vorzugsweise enthält die Hordenstruktur (6) 2 bis 20, besonders bevorzugt 2 bis 10 und ganz besonders bevorzugt 2 bis 6 Katalysatorbetten, die vorzugsweise axial (z. B. in einem Schachtofen) oder radial (z. B. in den Ringspalten von zentrisch ineinander gestellte zylindrischen Gitterrosten) angeordnet sind und vom Reaktionsgas in entsprechender Weise axial oder radial durchströmt werden;
die Hordenstruktur kann z. B. in einem einzigen Reaktor oder in einer Hintereinanderschaltung von Reaktoren verwirklicht sein;
die Katalysatorbetten sind vorzugsweise so mit Katalysatoren (z. B. jenen der DE-A 19937107, insbesondere den beispielhaft ausgeführten) beschickt, dass beim Durchströmen des Katalysatorbetts mit einem Reaktionsgas, dessen Zusammensetzung eine Konkurrenzreaktion zwischen heterogen katalysierter Wasserstoffverbrennung und heterogen katalysierter Propandehydrierung zulässt, in Strömungsrichtung zunächst die heterogen katalysierte Wasserstoffverbrennung schneller verläuft;
vorzugsweise wird dem dem ersten Abschnitt der Reaktionszone A zugeführten Reaktionsgasausgangsgemisch weder externer molekularer Sauerstoff, noch externer molekularer Wasserstoff, noch externer molekularer Wasserdampf zugeführt (unter einem an irgendeiner Stelle der Reaktionszone A (oder einer anderen Reaktions-/Trennzone des erfindungsgemäßen Verfahrens) extern zugeführten Gas soll in dieser Schrift generell Gas verstanden werden, das weder in der Reaktionszone A (noch in einer der anderen Reaktions-/Trennzonen des erfindungsgemäßen Verfahrens) selbst gebildet wird, noch in Kreisfahrweise aus der Reaktionszone A herauskommend unmittelbar und/oder mittelbar in die Reaktionszone A rückgeführt wird);
vorzugsweise wird dem Reaktionsgas(ausgangs)gemisch (des ersten Abschnitts der Reaktionszone A), nachdem es in Strömungsrichtung das erste Katalysatorbett durchlaufen hat, vor jedem Durchströmen eines dem ersten Katalysatorbett in Strömungsrichtung nachfolgenden Katalysatorbett des ersten Abschnitts der Reaktionszone A in begrenztem Umfang ein molekularen Sauerstoff enthaltendes Gas zugesetzt, wobei die dem Reaktionsgasgemisch so jeweils zugeführte Sauerstoffmenge jeweils so begrenzt wird, dass der resultierende Sauerstoffgehalt des Reaktionsgasgemischs, bezogen auf den in ihm enthaltenen Wasserstoff, 0,5 bis 50 bzw. bis 30 Vol.-%, vorzugsweise 10 bis 25 Vol.-% beträgt;
als ein solches Sauerstoff enthaltendes Gas wird vorzugsweise Luft (7) verwendet; es kann jedoch auch reiner Sauerstoff oder ein Gemisch aus Sauerstoff und Inertgasen wie CO₂, N₂ und/oder Edelgasen oder Stickoxid verwendet werden;
nach Zusatz des Sauerstoff enthaltenden Gases durchläuft das Reaktionsgas vorzugsweise noch einen statischen Mischer (8), bevor es in das relevante Katalysatorbett eintritt;
die Temperaturen innerhalb des ersten Abschnitts der Reaktionszone A werden vorzugsweise auf 400 bis 700°C gehalten;
der Arbeitsdruck innerhalb des ersten Abschnitts der Reaktionszone A liegt vorzugsweise bei 0,5 bis 10 bar;
das dem ersten Abschnitt der Reaktionszone A zugeführte Reaktionsgasausgangsgemisch (9) (seine Temperatur beträgt vorzugsweise 400 bis 700°C (z. B. 560°C), sein Druck liegt bevorzugt bei 0,5 bis 10 bar) weist vorzugsweise nachfolgende Gehalte auf:
10 bis 35 (z. B. 19 oder 28) Vol.-% Propan,
3 bis 10 (z. B. 7,5 oder 5,5) Vol.-% Propen,
5 bis 20 (z. B. 9 oder 12) Vol.-% Wasserdampf,
0,01 bis 6 (z. B. 3,9 oder 4,3) Vol.-% molekularer Wasserstoff und
0 bis 0,02 Vol.-% molekularer Sauerstoff;
die Belastung der Katalysatorbetten mit Reaktionsgas beträgt häufig 250 bis 5000 h^–1 (bei Hochlastfahrweise auch bis 40000 h^–1) vorzugsweise 10000 bis 25000 Nl/l·h, besonders bevorzugt 15000 bis 20000 Nl/l·h;
um das dem ersten Abschnitt der Reaktionszone A zugeführte Frischpropan auf die Reaktionstemperatur zu erwärmen, wird dieses vorteilhaft in indirektem Wärmeaustausch (in einem Gaskühler üblicher Bauart) zum aus dem zweiten Abschnitt der Reaktionszone A in die Trennzone A geführten heißen Produktgasgemisch A geführt;
außer den in dieser Schrift bereits genannten, werden dem ersten Abschnitt der Reaktionszone A vorzugsweise keine weiteren externen Gase zugeführt;
das im ersten Abschnitt der Reaktionszone A gebildete Produktgasgemisch A* weist vorzugsweise nachfolgende Gehalte auf:
8 bis 32 Vol.-% (z. B. 16 oder 24 Vol.-%) Propan,
3 bis 12 Vol.-% (z. B. 9 oder 6 Vol.-%) Propen,
5 bis 23 Vol.-% (z. B. 11 oder 15 Vol.-%) Wasserdampf,
0,1 bis 6 Vol.-% (z. B. 1 oder 3 Vol.-%) molekularer Wasserstoff, und
0 bis 0,05 Vol.-% molekularer Sauerstoff;
es weist vorzugsweise eine Temperatur von 500 bis 750°C (z. B. 580°C) und vorteilhaft einen Arbeitsdruck von 0,05 bis 10 bar auf.

• II. dem Produktgemisch A* anschließend molekularen Sauerstoff, Propan und gegebenenfalls (in der Reaktionszone B) nicht umsetztes Propylen enthaltendes (Haupt- und/oder Neben- bzw. Gesamt-)Restgas aus der Trennzone B zuführt (erfindungsgemäß vorteilhaft wenigstens die Hälfte, bevorzugt wenigstens drei Vierteil und ganz besonders bevorzugt die Gesamtmenge des vorgenannten in der Trennzone B anfallenden Restgases) und das dabei resultierende Reaktionsgasgemisch A* dem zweiten Abschnitt der Reaktionszone A zuführt;

das dem Produktgasgemisch A* zugeführte Restgas enthält in der Regel 0,5 Vol.-% bis 10 Vol.-%, häufig 1 bis 8 Vol.-% und bevorzugt 2 bis 5 Vol.-% an molekularem Sauerstoff;
die Gehalte des dem Produktgasgemisch A* zugeführten (Haupt- und/oder Neben- bzw. Gesamt-)Restgases sind vorzugsweise wie folgt:
10 bis 40 Vol.-% (z. B. 15 oder 32 Vol.-%) Propan,
0 bis 1 Vol.-% Propen,
> 0 bis 5 Vol.-% (z. B. 3 oder 4 Vol.-%) molekularer Sauerstoff,
1 bis 10 Vol.-% (z. B. 2 Vol.-%) Wasserdampf, und
0 bis 0,5 Vol.-% (z. B. 0 bis 0,3 oder 0 bis 0,1 Vol.-%) molekularer Wasserstoff;
vorzugsweise wird dem Produktgasgemisch A* das Restgas in solchen Mengen zugeführt, dass der Sauerstoffgehalt des resultierenden Reaktionsgasgemischs A*, bezogen auf den in ihm enthaltenen Wasserstoff, 15 bis 50 Vol.-%, vorteilhaft 25 bis 50 Vol.-%, und besonders vorteilhaft 30 oder 40 bis 50 Vol.-% beträgt;
das Reaktionsgasgemisch A* (12) weist vorzugsweise nachfolgende Gehalte auf:
10 bis 38 Vol.-% (z. B. 16 oder 28 Vol.-%) Propan,
3 bis 10 Vol.-% (z. B. 5 Vol.-%) Propen,
0 bis 0,05 Vol.-% molekularer Sauerstoff,
3 bis 20 Vol.-% (z. B. 7 oder 12 Vol.-%) Wasserdampf, und
0,1 bis 6 Vol.-% (z. B. 4 Vol.-%) molekularer Wasserstoff.
die Temperatur des Reaktionsgasgemischs A* beträgt vorteilhaft 300 bis 600°C (z. B. 400 bis 500°C), der Druck des Reaktionsgasgemischs A* beträgt vorteilhaft 0,5 bzw. 0,6 bis 12 bar;
bevorzugt führt man dem Produktgasgemisch A* (11) das (Haupt- und/oder Neben- bzw. Gesamt-)Restgas aus der Trennzone B nach dem Prinzip einer mit diesem Restgas als Treibstrahl betriebenen Strahlpumpe zu, wobei die Förderrichtung des durch eine Treibdüse (1) über eine Mischstrecke (2) und einen Diffusor (3) entspannten Treibstrahls in den zweiten Abschnitt der Reaktionszone A (13) sowie die Saugrichtung (die Saugwirkung) des Saugstutzens (4) in Richtung des ersten Abschnitts der Reaktionszone A weist und die Verbindung „Saugstutzen-Mischstrecke-Diffusor" die alleinige Verbindung zwischen dem ersten und dem zweiten Abschnitt der Reaktionszone A bildet;
vorzugsweise weist der Treibstrahl eine Temperatur von 400 bzw. 500 bis 600°C und einen Druck von 3 bis 6 bar (vorzugsweise 4 bis 5 bar) auf;
vorzugsweise wird das aus der Trennzone B in die Reaktionszone A geführte Restgas (15) durch indirekten Wärmetausch (14) mit in die Trennzone A geführtem heißem Produktgasgemisch A (16) auf diese Temperatur erwärmt (vorzugsweise nachdem das Frischpropan bereits mit diesem Produktgasgemisch A erwärmt worden ist);
vorzugsweise wird das aus der Trennzone B in die Reaktionszone A geführte Restgas mittels eines Radial- bzw. Axialverdichters auf vorgenannten Druck verdichtet;
vorzugsweise enthält das Reaktionsgasgemisch A* außer den in dieser Schrift genannten Gasströmen keine weiteren Gasströme zugesetzt;

• III. im zweiten Abschnitt der Reaktionszone A unter Ausbildung des Propan und Propylen enthaltenden Produktgasgemisch A im Reaktionsgasgemisch A* enthaltenen molekularen Sauerstoff mit darin enthaltenem molekularem Wasserstoff heterogen katalysiert zu Wasser verbrennt (vorteilhaft wenigstens 25 mol-%, bevorzugt wenigstens 50 mol-%, besonders bevorzugt wenigstens 75 mol-%, besser wenigstens 90 mol-% und am Besten wenigstens 95 mol-% der im Reaktionsgasgemisch A* enthaltenen Menge an molekularem Sauerstoff) und gegebenenfalls im Reaktionsgasgemisch A* enthaltenes Propan heterogen katalysiert zu Propylen dehydriert (in der Regel weniger als 40 mol-% bzw. weniger als 30 mol-%, meist weniger als 20 mol-%, häufig weniger als 10 mol-%, und oft weniger als 5 mol-% des im Reaktionsgasgemisch A* enthaltenen Propan);

dabei wird der zweite Abschnitt der Reaktionszone A im wesentlichen adiabat betrieben (d.h., über außerhalb des zweiten Abschnitts der Reaktionszone A befindliche Wärmeträger wird dem zweiten Abschnitt der Reaktionszone A im wesentlichen weder (gezielt) Wärme zugeführt noch entnommen);
bezogen auf einmaligen Durchgang des Reaktionsgasgemischs A* durch den zweiten Abschnitt der Reaktionszone A ist die Bruttowärmetönung über den zweiten Abschnitt der Reaktionszone A exotherm (um diese positive Wärmetönung zu gewährleisten, kann es insbesondere im Fall von vergleichsweise geringen Dehydrierumsätzen im ersten Abschnitt der Reaktionszone zweckmäßig sein, den Wasserstoffgehalt des Reaktionsgasgemischs A* durch Zufuhr von externem Wasserstoff zu erhöhen; anwendungstechnisch zweckmäßig wird man diese Zufuhr von molekularem Wasserstoff vorteilhaft bereits ins Produktgasgemisch A* hinein vornehmen, sie kann aber z. B. auch in das rückgeführte Restgas hinein erfolgen (19) oder in das resultierende Reaktionsgasgemisch A*);
außer den in dieser Schrift bereits genannten, werden dem zweiten Abschnitt der Reaktionszone A vorzugsweise keine weiteren externen Gase zugeführt;
die Temperaturen innerhalb des zweiten Abschnitts der Reaktionszone A werden vorzugsweise auf 400 bis 750°C (bzw. 500 bis 700°C) gehalten;
der Arbeitsdruck innerhalb des zweiten Abschnitts der Reaktionszone A liegt vorzugsweise bei 0,5 bis 12 bar;
für den zweiten Abschnitt der Reaktionszone A sind z. B. diejenigen Katalysatoren geeignet, die in dieser Schrift auch für die heterogen katalysierte Dehydrierung im ersten Abschnitt der Reaktionszone A als geeignet empfohlen werden, da sie wie eingangs dieser Schrift bereits erwähnt, in der Regel auch die Verbrennung von molekularem Wasserstoff zu katalysieren vermögen (dies trifft insbesondere auf die Katalysatoren der DE-A 19937107 (insbesondere auf die dort beispielhaft aufgeführten) zu, weshalb diese Katalysatoren ganz besonders bevorzugt in beiden Abschnitten der Reaktionszone A zum Einsatz kommen; in einer Konkurrenzsituation zwischen heterogen katalysierter Propandehydrierung und heterogen katalysierter Wasserstoffverbrennung verläuft letztere an den genannten Katalysatoren in der Regel wesentlich schneller und dominiert die erstere); selbstredend kommen für den zweiten Abschnitt der Reaktionszone A aber auch solche Katalysatoren in Betracht, die spezifisch für die selektive Verbrennung von molekularem Wasserstoff maßgeschneidert sind; solche Katalysatoren sind z. B. jene der Schriften US-A 4788371, US-A 4886928, US-A 5430203, US-A 5530171, US-A 5527979 und US-A 5563314;
als Reaktoren kommen für den zweiten Abschnitt der Reaktionszone A grundsätzlich dieselben in Betracht, wie für den ersten Abschnitt der Reaktionszone A;
vorzugsweise enthält der Reaktor für den zweiten Abschnitt der Reaktionszone A keine Vielzahl von Katalysatorbetten, sondern nur ein Katalysatorbett;
anwendungstechnisch zweckmäßig ist dieses eine Katalysatorbett in einem einfachen Schachtofen untergebracht und wird axial oder radial durchströmt;
die Belastung des vorgenannten Katalysatorbetts mit Reaktionsgasgemisch A* beträgt in der Regel 500 bis 80000 Nl/l·h, vorzugsweise 10000 bis 50000 Nl/l·h, besonders bevorzugt 20000 bis 40000 Nl/l·h (bei nur geringen Dehydrierumsätzen im zweiten Abschnitt der Reaktionszone A und vorwiegend ausschließlicher Verbrennung von molekularem Sauerstoff mit molekularem Wasserstoff kann die vorgenannte Belastung auch bis zu 200000 Nl/l·h betragen);
mit Vorteil sind der erste und der zweite Abschnitt der Reaktionszone A in ein und demselben Reaktor untergebracht;
anwendungstechnisch vorteilhaft ist der zweite Abschnitt der Reaktionszone A eine mit Katalysator beschickte Horde in einem Hordenreaktor, dessen übrige Horden dem ersten Abschnitt der Reaktionszone A zugeordnet sind;
das im zweiten Abschnitt der Reaktionszone A gebildete Produktgasgemisch A weist vorzugsweise nachfolgende Gehalte auf:
7 bis 30 Vol.-% (z. B. 12 oder 24 Vol.-%) Propan,
3 bis 12 Vol.-% (z. B. 6 oder 8 Vol.-%) Propen,
0 bis 0,05 Vol.-% molekularer Sauerstoff,
5 bis 22 Vol.-% (z. B. 10 oder 13 Vol.-%) Wasserdampf und
0,1 bis 6 Vol.-% (z. B. 4 Vol.-%) molekularer Wasserstoff.
die Temperatur des Produktgasgemischs A beträgt vorteilhaft 400 bis 700°C; der Druck des Reaktionsgasgemischs beträgt vorteilhaft 0,4 bis 10 bar und liegt vorzugsweise gleichzeitig oberhalb des Drucks des Produktgasgemischs A*;

• IV. das im zweiten Abschnitt der Reaktionszone A ausgebildete Produktgasgemisch A in zwei Teilmengen identischer Zusammensetzung aufteilt und die eine der beiden Teilmengen (10) als ein weiterer Propan enthaltender Zufuhrstrom (als Dehydrierkreisgas) in den ersten Abschnitt der Reaktionszone A rückführt (bevorzugt als Bestandteil des Reaktionsgasausgangsgemischs des ersten Abschnitts der Reaktionszone A und vorzugsweise in natürlichem Umlauf dem zum Druck am Ausgang des ersten Abschnitts der Reaktionszone A bestehenden Druckgefälle folgend; die Vermischung mit Frischpropan erfolgt vorzugsweise mittels statischem Mischer vor dem in Strömungsrichtung befindlichen ersten Katalysatorbett des ersten Abschnitts der Reaktionszone A) und die andere Teilmenge (18) an Produktgasgemisch A (vorzugsweise die gesamte andere Teilmenge) in der ersten Trennzone A erfindungsgemäß weiterbehandelt;

vorzugsweise führt man die andere Teilmenge (18) an Produktgasgemisch A über einen indirekten Wärmeaustauscher (14) in die Trennzone A, um in selbigem das Frischpropan und/oder das Restgas aus der Trennzone B (auf die gewünschte Temperatur) zu erwärmen.A very particularly preferred embodiment of the reaction zone A of the method according to the invention is therefore as follows and is shown schematically in FIG 1 This document, to which all numerical addresses refer, is shown graphically:
The reaction zone A ( 17 ) is designed and operated as already described as a combination of the above-described "loop variant" and "two-section variant", wherein additionally the following specifications are met:

• I. the first section ( 0 ) feeds to reaction zone A at least one supply stream containing gaseous propane, the fresh propane ( 5 ), and in this first section of the reaction zone A subjects the propane fed to it to a heterogeneously catalyzed (partial) dehydrogenation to obtain a propane, propylene and molecular hydrogen-containing product gas mixture A *, at least 5 mol%, or at least 10 mol%, preferably at least 15 mol%, or at least 20 mol%, more preferably at least 25 mol%, or at least 30 mol%, most preferably at least 35 mol%, or at least 40 mol%, and more preferably at least 45 mole%, or at least 50 mole% (but usually less than 70 mole%, or less than 60 mole%, or ≤50 mole%) of the first portion of reaction zone A (total Propane in this first section have been reacted dehydrating (in loop mode in the reaction zone A are preferred process variants in turn, those in which the aforementioned turnover numbers Z _U each by Z _U ^Kr = (Z _U / 1 + K GV) are replaced);

In this case, the first section of the reaction zone A is operated substantially adiabatically (ie about outside the first portion of the reaction zone A heat transfer medium is the first portion of the reaction zone A substantially neither (targeted) heat supplied nor removed);
based on a single pass of the feed gas mixture of the first section of reaction zone A through the first section of reaction zone A, the gross thermal tint over the first section of reaction zone A is endothermic (negative) to autothermal (essentially zero);
Preferably, the first portion of the reaction zone A is carried out in Hordenstruktur;
preferably contains the Hordenstruktur ( 6 ) 2 to 20, particularly preferably 2 to 10 and very particularly preferably 2 to 6 catalyst beds, which are preferably arranged axially (eg in a shaft furnace) or radially (eg in the annular gaps of centrically interlocked cylindrical gratings) and flowed through by the reaction gas in a corresponding manner axially or radially;
the horde structure can z. B. be implemented in a single reactor or in a series connection of reactors;
the catalyst beds are preferably charged with catalysts (for example those of DE-A 19937107, in particular those exemplified) that, when flowing through the catalyst bed with a reaction gas whose composition allows a competing reaction between heterogeneously catalyzed hydrogen combustion and heterogeneously catalyzed propane dehydrogenation Flow direction initially the heterogeneously catalyzed hydrogen combustion proceeds faster;
Preferably, the reaction gas starting mixture fed to the first section of reaction zone A is supplied with neither external molecular oxygen nor external molecular hydrogen nor external molecular water vapor (under a gas supplied externally at any point in reaction zone A (or another reaction / separation zone of the process according to the invention) In this document, gas should generally be understood which is not formed either in reaction zone A (or in one of the other reaction / separation zones of the process according to the invention), nor in recycle mode coming out of reaction zone A directly and / or indirectly into reaction zone A. is returned);
Preferably, the reaction gas (outlet) mixture (the first portion of the reaction zone A), after having passed through the first catalyst bed in the flow direction, before each flowing through a first catalyst bed in the flow direction subsequent catalyst bed of the first portion of the reaction zone A to a limited extent added a molecular oxygen-containing gas, the reaction gas mixture so the respectively supplied amount of oxygen is respectively limited so that the resulting oxygen content of the reaction gas mixture, based on the hydrogen contained in it 0.5 to 50 or 30% by volume, preferably 10 to 25% by volume;
as such an oxygen-containing gas is preferably air ( 7 ) used; however, it is also possible to use pure oxygen or a mixture of oxygen and inert gases such as CO ₂ , N ₂ and / or noble gases or nitrogen oxide;
after addition of the oxygen-containing gas, the reaction gas preferably also passes through a static mixer ( 8th ) before entering the relevant catalyst bed;
the temperatures within the first section of the reaction zone A are preferably maintained at 400 to 700 ° C;
the working pressure within the first section of the reaction zone A is preferably from 0.5 to 10 bar;
the reaction gas starting mixture fed to the first section of the reaction zone A ( 9 ) (its temperature is preferably 400 to 700 ° C (eg 560 ° C), its pressure is preferably 0.5 to 10 bar) preferably has the following contents:
10 to 35 (eg, 19 or 28)% by volume of propane,
3 to 10 (eg, 7.5 or 5.5)% by volume of propene,
From 5 to 20 (eg 9 or 12)% by volume of water vapor,
0.01 to 6 (eg 3.9 or 4.3)% by volume of molecular hydrogen and
0 to 0.02% by volume of molecular oxygen;
the loading of the catalyst beds with the reaction gas is frequently 250 to 5000 h ^-1 (also up to 40,000 h ^-1 in the case of high-load operation), preferably 10000 to 25,000 Nl / l · h, particularly preferably 15,000 to 20,000 Nl / l · h;
in order to heat the fresh propane supplied to the first section of the reaction zone A to the reaction temperature, this is advantageously conducted in indirect heat exchange (in a gas cooler of conventional design) to the hot product gas mixture A conducted from the second section of the reaction zone A into the separation zone A;
Apart from those already mentioned in this document, the first section of the reaction zone A preferably no further external gases are supplied;
the product gas mixture A * formed in the first section of the reaction zone A preferably has the following contents:
8 to 32% by volume (eg 16 or 24% by volume) of propane,
3 to 12% by volume (eg 9 or 6% by volume) of propene,
5 to 23% by volume (eg 11 or 15% by volume) of water vapor,
0.1 to 6 vol% (eg, 1 or 3 vol%) of molecular hydrogen, and
0 to 0.05% by volume of molecular oxygen;
it preferably has a temperature of 500 to 750 ° C (eg 580 ° C) and advantageously a working pressure of 0.05 to 10 bar.

• II. The product mixture A * then molecular oxygen, propane and optionally (in the reaction zone B) unreacted propylene containing (main and / or secondary or total) residual gas from the separation zone B feeds (according to the invention advantageously at least half, preferably at least three-fourths and most preferably the total amount of the abovementioned residual gas arising in the separation zone B) and the resulting reaction gas mixture A * is fed to the second section of the reaction zone A;

the residual gas fed to the product gas mixture A * generally contains from 0.5% by volume to 10% by volume, frequently from 1 to 8% by volume and preferably from 2 to 5% by volume, of molecular oxygen;
the contents of the (main and / or secondary or total) residual gas supplied to the product gas mixture A * are preferably as follows:
10 to 40% by volume (eg 15 or 32% by volume) of propane,
0 to 1% by volume of propene,
> 0 to 5% by volume (eg 3 or 4% by volume) of molecular oxygen,
1 to 10% by volume (eg 2% by volume) of water vapor, and
0 to 0.5 vol% (eg 0 to 0.3 or 0 to 0.1 vol%) of molecular hydrogen;
Preferably, the residual gas is supplied to the product gas mixture A * in amounts such that the oxygen content of the resulting reaction gas mixture A *, based on the hydrogen contained in it, 15 to 50 vol .-%, advantageously 25 to 50 vol .-%, and particularly advantageous Is 30 or 40 to 50% by volume;
the reaction gas mixture A * ( 12 ) preferably has the following contents:
10 to 38% by volume (eg 16 or 28% by volume) of propane,
From 3 to 10% by volume (for example 5% by volume) of propene,
0 to 0.05% by volume of molecular oxygen,
From 3 to 20% by volume (eg 7 or 12% by volume) of water vapor, and
0.1 to 6 vol% (eg, 4 vol%) of molecular hydrogen.
the temperature of the reaction gas mixture A * is advantageously 300 to 600 ° C. (eg 400 to 500 ° C.), the pressure of the reaction gas mixture A * is advantageously 0.5 or 0.6 to 12 bar;
Preferably, the product gas mixture A * ( 11 ) the (main and / or secondary or total) residual gas from the separation zone B according to the principle of a driven with this residual gas as a jet jet jet pump, wherein the conveying direction of the by a motive nozzle ( 1 ) via a mixing section ( 2 ) and a diffuser ( 3 ) in the second section of the reaction zone A ( 13 ) as well as the suction direction (the suction effect) of the suction nozzle ( 4 ) in the direction of the first section of the reaction zone A and the connection "suction port mixing section diffuser" forms the sole connection between the first and the second section of the reaction zone A;
Preferably, the propulsion jet has a temperature of 400 or 500 to 600 ° C and a pressure of 3 to 6 bar (preferably 4 to 5 bar);
Preferably, the outgoing from the separation zone B in the reaction zone A residual gas ( 15 ) by indirect heat exchange ( 14 ) with hot product gas mixture A conducted into the separation zone A ( 16 ) is heated to this temperature (preferably after the fresh propane has already been heated with this product gas mixture A);
Preferably, the guided from the separation zone B in the reaction zone A residual gas is compressed by means of a radial or axial compressor to the aforementioned pressure;
Preferably, the reaction gas mixture A * contains no further gas streams added except for the gas streams mentioned in this document;

• III. in the second section of the reaction zone A with formation of the propane and propylene-containing product gas mixture A in the reaction gas mixture A * contained molecular oxygen containing molecular hydrogen heterogeneously catalyzed to water (advantageously at least 25 mol%, preferably at least 50 mol%, more preferably at least 75 mol%, more preferably at least 90 mol% and most preferably at least 95 mol% of the amount of molecular oxygen contained in the reaction gas mixture A *) and optionally in the reaction gas mixture A * propane heterogeneously catalyzed dehydrogenated to propylene (usually less than 40 mol% or less than 30 mol%, usually less than 20 mol%, often less than 10 mol%, and often less than 5 mol% of the propane contained in the reaction gas mixture A *);

the second section of the reaction zone A is operated essentially adiabatically (ie heat transfer medium located outside the second section of the reaction zone A is supplied with heat to the second section of the reaction zone A substantially neither (deliberately));
based on a single pass of the reaction gas mixture A * through the second section of the reaction zone A, the gross thermal tinting over the second section of the reaction zone A is exothermic (to ensure this positive heat of reaction, it may be useful in the first section of the reaction zone, especially in the case of comparatively low dehydrogenation conversions to increase the hydrogen content of the reaction gas mixture A * by supplying external hydrogen, this feed of molecular hydrogen will advantageously already be made into the product gas mixture A *, but it can also be carried into the recycled residual gas, for example. 19 ) or in the resulting reaction gas mixture A *);
Apart from those already mentioned in this document, the second section of the reaction zone A preferably no further external gases are supplied;
the temperatures within the second section of the reaction zone A are preferably maintained at 400 to 750 ° C (or 500 to 700 ° C);
the working pressure within the second section of the reaction zone A is preferably 0.5 to 12 bar;
for the second section of the reaction zone A are z. For example, those catalysts are suitable, which are recommended in this document as suitable for the heterogeneously catalyzed dehydrogenation in the first section of the reaction zone A, since they as mentioned at the beginning of this document, usually catalyze the combustion of molecular hydrogen assets (this applies in particular to the catalysts of DE-A 19937107 (in particular to those exemplified there), which is why these catalysts are very particularly preferably used in both sections of reaction zone A. In a competition situation between heterogeneously catalyzed propane dehydrogenation and heterogeneously catalyzed hydrogen combustion, the latter proceeds the catalysts mentioned usually much faster and dominates the former); Of course, suitable for the second section of the reaction zone A but also those catalysts are specifically tailored for the selective combustion of molecular hydrogen; Such catalysts are for. U.S. Patent Nos. 4,788,371, 4,886,928, 5,430,203, 5,530,171, 5,527,979 and 5,563,314;
as reactors, the same are in principle possible for the second section of the reaction zone A as for the first section of the reaction zone A;
Preferably, the reactor for the second section of the reaction zone A contains no plurality of catalyst beds, but only one catalyst bed;
In terms of application technology, this catalyst bed is accommodated in a simple shaft furnace and is flowed through axially or radially;
the load of the aforementioned catalyst bed with reaction gas mixture A * is generally 500 to 80,000 Nl / l.h, preferably 10,000 to 50,000 Nl / l.h, more preferably 20,000 to 40,000 Nl / l.h (with only low dehydrogenation conversions in the second section the reaction zone A and predominantly exclusive combustion of molecular oxygen with molecular hydrogen, the aforementioned load may also be up to 200,000 Nl / l · h);
Advantageously, the first and second sections of reaction zone A are housed in one and the same reactor;
In terms of application technology, the second section of the reaction zone A is a catalyst-loaded horde in a tray reactor whose remaining hordes are assigned to the first section of the reaction zone A;
the product gas mixture A formed in the second section of the reaction zone A preferably has the following contents:
7 to 30% by volume (eg 12 or 24% by volume) of propane,
3 to 12% by volume (eg 6 or 8% by volume) of propene,
0 to 0.05% by volume of molecular oxygen,
5 to 22% by volume (eg 10 or 13% by volume) of water vapor and
0.1 to 6 vol% (eg, 4 vol%) of molecular hydrogen.
the temperature of the product gas mixture A is advantageously 400 to 700 ° C; the pressure of the reaction gas mixture is advantageously from 0.4 to 10 bar and is preferably at the same time above the pressure of the product gas mixture A *;

• IV. The formed in the second section of the reaction zone A product gas mixture A is divided into two subsets of identical composition and one of the two subsets ( 10 ) as a further propane-containing feed stream (as Dehydrierkreisgas) in the first portion of the reaction zone A (preferably as part of the starting reaction gas mixture of the first portion of the reaction zone A and preferably in natural circulation following the pressure at the exit of the first portion of the reaction zone A pressure gradient mixing with fresh propane preferably takes place by means of a static mixer upstream of the first catalyst bed of the first section of the reaction zone A) located in the direction of flow, and the other subset ( 18 ) to product gas mixture A (preferably the entire other subset) in the first separation zone A further treated according to the invention;

Preferably, you lead the other subset ( 18 ) to product gas mixture A via an indirect heat exchanger ( 14 ) in the separation zone A, to heat in the same the fresh propane and / or the residual gas from the separation zone B (to the desired temperature).

The product gas mixture A formed in the process according to the invention in the heterogeneously catalyzed propane dehydrogenation in the reaction zone A generally contains propane, propene, molecular hydrogen, N ₂ , H ₂ O, methane, ethane, ethylene, butene-1, other butenes (eg B. iso-butene) and other C ₄ hydrocarbons (n-butane, isobutane, butadiene, etc.), CO and CO ₂ , but usually also oxygenates such as alcohols, aldehydes and carboxylic acids (usually with ≤ 9 C). atoms). Furthermore, in small quantities from the residual gas originating ingredients may be included, such. B. minor components which are formed in the reaction zone B, but also in the separation zones A and / or B used absorbent. It will usually be at a pressure of 0.3 to 10 bar and often have a temperature of 400 to 650 or to 500 ° C, in favorable cases from 450 to 500 ° C.

While the EP-A 117 146, DE-A 3 313 573 and US-A 3 161 670, that formed in heterogeneously catalyzed propane dehydrogenation Product gas mixture A as such for charging the reaction zone B, it is necessary according to the invention, from the product gas mixture A prior to its further use for charging the reaction zone B as Nebenkomponentenauslass at least a subset of therein contained, other than propane and propylene, components, separate. The requirements of DE-A 10211275 must be observed. contains the product gas mixture A still significant amounts of molecular Hydrogen, the aforementioned separation is usually at least associated with a partial separation of the hydrogen or you will advantageously carry out such a hydrogen separation in advance.

The latter can z. B. be done by reacting the product gas mixture A, optionally after having previously cooled in an indirect heat exchanger (Expediently will, as already mentioned, in the reaction zone A, the heat removed for heating a process for the invention required Feed gas used) one, usually to a pipe (but it is also possible Plate or a winding module) designed, the membrane conducts only for the permeable to molecular hydrogen is. The thus separated molecular hydrogen can partially if necessary into the heterogeneously catalyzed dehydrogenation of propane (i.e., in the reaction zone A) recycled and / or recycled become. For example, it can be burned in fuel cells.

alternative can be a partial or complete Hydrogen separation also by partial condensation, adsorption and / or rectification (preferably under pressure). Partial or full separation of molecular hydrogen from product gas mixture A can in the method according to the invention also by outside Reaction zone A selective (eg, heterogeneously catalyzed) Combustion thereof be made with molecular oxygen. The resulting water of reaction can be either partially or completely be separated or left in the gas mixture, as it is in the reaction zone B as an inert diluent gas to act. In this regard, suitable catalysts disclose, for example, US-A 4,788 371, US Pat. No. 4,886,928, US Pat. No. 5,430,209, US Pat. No. 5,530,171, US Pat. No. 5,527,979 and US-A 5,563,314.

According to the invention is appropriate at least 10 mol%, or at least 25 mol%, often at least 35 mol%, or at least 50 mol%, often at least 75 mol% and often the total amount of the product gas mixture led out of the reaction zone A. Pre-and / or co-separate molecular hydrogen contained in A, before the remaining product gas mixture A 'to charge the reaction zone B is used. However, it is a very advantageous feature the most recently described, most preferred Operation of the reaction zone A, namely the combination of "loop variant" and "two-section variant", with adiabatic endothermic Design of the first section and adiabatic exothermic design of the second section (each of the reaction zone A) that there the product gas mixture A already largely free of molecular Hydrogen is produced, the molecular hydrogen formed in reaction zone A. is still used almost quantitatively advantageous in the reaction zone A. and the water vapor formed as part of the in the first Section of the reaction zone A recycled partial amount the product gas mixture A also an advantageous use supplied becomes. If necessary, from the product gas mixture A before its further processing in the reaction zone B also possibly contained water in part or completely separated (eg, condensing). Of course, at Need in the context of the separation of molecular hydrogen and / or Water vapor also a separation of other, propane and propylene various, components of the product gas mixture A made become.

A easy way this is z. B. therein, preferably cooled (preferably to temperatures from 10 to 100 or 70 ° C), Product gas mixture A, z. B. at a pressure of 0.1 to 50 bar, preferably 5 to 15 bar and a temperature of z. B. 0 to 100 ° C, preferably 20 to 40 ° C, with a (preferably high boiling) organic solvent (preferably a hydrophobic) in which propane and propylene (opposite to the other constituents of the product gas mixture A expediently preferred) be brought into contact (eg by simple Passing). By subsequent desorption, rectification and / or Stripping with a reference the reaction zone B is inert and / or in this reaction zone needed as a reactant Gas (eg, air or another mixture of molecular oxygen and inert gas), the propane and propylene in the mixture in purified Form recovered and used to charge the reaction zone B (preferably The procedure is as described in Comparative Example 1) The optionally Molecular hydrogen-containing offgas of the absorption can be z. B. subject to membrane separation again and then, if necessary, the separated hydrogen in the heterogeneously catalyzed propane dehydrogenation in the reaction zone A use.

Indeed is the C3 hydrocarbons / C4 hydrocarbons separation factor in the above separation method comparatively limited and for in DE-A 10245585 described needs often not sufficient.

When alternative for the separation step described above via absorption is for the purposes according to the invention therefore often a pressure swing adsorption or a pressure rectification preferred.

In principle, all absorbents which are capable of absorbing propane and propylene are suitable as absorbents for the absorptive separation described above. In the Absorptionsmit It is preferably an organic solvent, which is preferably hydrophobic and / or high-boiling. Advantageously, this solvent has a boiling point (at a normal pressure of 1 atm) of at least 120 ° C, preferably of at least 180 ° C, preferably from 200 to 350 ° C, especially from 250 to 300 ° C, more preferably from 260 to 290 ° C. Conveniently, the flash point (at a normal pressure of 1 bar) is above 110 ° C. Generally suitable as absorbent relatively non-polar organic solvents, such as aliphatic hydrocarbons, which preferably contain no polar group acting outward, but also aromatic hydrocarbons. In general, it is desirable that the absorbent has the highest possible boiling point with the highest possible solubility for propane and propylene. Examples which may be mentioned as absorbents are aliphatic hydrocarbons, for example C ₈ -C ₂₀ alkanes or alkenes, or aromatic hydrocarbons, for example, middle oil fractions from the Paraffindestillation or ethers with bulky (sterically demanding) groups on the O atom, or mixtures thereof, wherein this may be added a polar solvent, such as, for example, the 1,2-dimethyl phthalate disclosed in DE-A 43 08 087. Also suitable are esters of benzoic acid and phthalic acid with straight-chain alkanols containing 1 to 8 carbon atoms, such as n-butyl benzoate, benzoic acid, ethyl benzoate, dimethyl phthalate, diethyl phthalate, and so-called heat transfer oils, such as diphenyl, diphenyl ether and mixtures of diphenyl and diphenyl ether or their chlorinated derivatives and triaryl alkenes, for example, 4-methyl-4'-benzyl-diphenylmethane and its isomers, 2-methyl-2'-benzyl-diphenylmethane, 2-methyl-4'-benzyldiphenylmethane and 4-methyl-2'-benzyl-diphenylmethane and mixtures of such isomers. A suitable absorbent is a solvent mixture of diphenyl and diphenyl ether, preferably in the azeotropic composition, in particular from about 25 wt .-% diphenyl (biphenyl) and about 75 wt .-% of diphenyl ether, for example the commercially available Diphyl ^® (eg. B of of Bayer Aktiengesellschaft). Frequently, this solvent mixture contains a solvent such as dimethyl phthalate in an amount of 0.1 to 25 wt .-%, based on the total solvent mixture added. Particularly suitable absorbents are also octanes, nonanes, decanes, undecanes, dodecanes, tridecanes, tetradecanes, pentadecanes, hexadecanes, heptadecans and octadecanes, with tetradecanes in particular having proven to be particularly suitable. It is favorable if the absorbent used on the one hand fulfills the above-mentioned boiling point, but on the other hand has a molecular weight which is not too high. Advantageously, the molecular weight of the absorbent is <300 g / mol. Also suitable are the paraffin oils having 8 to 16 carbon atoms described in DE-A 33 13 573. Examples of suitable commercial products are products sold by Haltermann, such as Halpasols i, such as Halpasol 250/340 i and Halpasol 250/275 i, as well as printing ink oils under the names PKWF and Printosol. Preference is given to aromatics-free commercial products, eg. B. those of the type PKWFaf.

The performance of absorption is not particularly limited. All methods and conditions customary to the person skilled in the art can be used. Preferably, the gas mixture at a pressure of 1 to 50 bar, preferably 2 to 20 bar, more preferably 5 to 15 bar, and a temperature of 0 to 100 ° C, in particular from 20 to 50 or 40 ° C, with the absorbent brought into contact. The absorption can be carried out both in columns and in quenching apparatuses. It can be done in cocurrent or in countercurrent (is preferred). Suitable absorption columns are, for example tray columns (having bubble and / or sieve trays), columns with structured packings (for example, sheet metal packings with a specific surface area of 100 to 1000, or up to 750 m ² / m ^3, for example Mellapak ^® 250 Y) and Packed columns (filled, for example, with Raschig fillers). However, it is also possible to use trickle and spray towers, graphite block absorbers, surface absorbers such as thick-film and thin-layer absorbers, dishwashers, cross-flow scrubbers and rotary scrubbers. In addition, it may be beneficial to let the absorption take place in a bubble column with and without internals.

The Separation of the propane and propylene from the absorbent can by stripping, flash evaporation and / or distillation respectively.

The separation of the propane and propylene from the absorbent is preferably carried out by stripping and / or desorption. The desorption can be carried out in a customary manner via a pressure and / or temperature change, preferably at a pressure of 0.1 to 10 bar, in particular 1 to 5 bar, more preferably 1 to 3 bar, and a temperature of 0 to 200 ° C, especially 20 to 100 ° C, more preferably 30 to 70 ° C, particularly preferably 30 to 50 ° C. A gas suitable for stripping is, for example, water vapor, but in particular oxygen / nitrogen mixtures, for example air, are preferred. When using air or oxygen / nitrogen mixtures in which the oxygen content is more than 10% by volume, it may be useful to add a gas before and / or during the stripping process, which reduces the explosion range. Particularly suitable for this are gases with a specific heat capacity ≥ 29 J / mol · K at 20 ° C, such as methane, ethane, propane, propene, benzene, methanol, ethanol, and ammonia, carbon dioxide, and water. C 4 -hydrocarbons are inventively preferred as sol However, additives should be avoided. Particularly suitable for the stripping are bubble columns with and without internals.

The Separation of the propane and propylene from the absorbent can also over a distillation or rectification are carried out, with those skilled in the art common Columns with packs, packing or corresponding internals can be used. Preferred conditions in the distillation or rectification are a pressure of 0.01 up to 5 bar, in particular 0.1 to 4 bar, more preferably 1 to 3 bar, and a temperature (in the bottom) of 50 to 300 ° C, especially 150 to 250 ° C.

One can be obtained by stripping from the absorbent gas mixture before its use for charging the reaction zone B one more supplied to further process stage be used to For example, the losses of co-stripped absorbent (eg separation in demisters and / or depth filters) and the reaction zone B at the same time before absorbing agent protect or to the separation effect between C3 / C4 hydrocarbons on to improve. Such separation of the absorbent can take place after all the process steps known in the art. One in the context of the method according to the invention preferred embodiment such a separation is z. As the quenching of the output current from the stripping device with water. In this case, the absorbent becomes washed out of this loaded output stream with water and the output current is simultaneously charged with water. This laundry or the quenching can z. B. at the top of a desorption column above a liquid collecting bottom by spraying carried out by water or in a separate apparatus.

to support of the separation effect can in the quench space, the quench surface enlarging internals be installed, as the expert of rectifications, absorptions and Desorptionsen ago known.

water is in this respect a preferred detergent, as it is in the following normally does not bother at least one partial zone. After this the water is the absorbent from the propane and propylene has washed out loaded feed stream, the water / absorbent mixture fed to a phase separation and the treated output stream immediately to the reaction zone B supplied become.

Either the substantially C3-free stripped absorbent as well the recovered in the phase separation Absorbents can for the Absorption purpose to be reused.

The in a separation which can be carried out as described by way of example, the propane and propene containing gas mixture can now in itself known manner for charging a heterogeneously catalyzed gas phase partial oxidation of propylene to acrolein and / or acrylic acid, such as it z. In WO 01/96270 and in DE-A 10245585, DE-A 10246119, DE-A 10313210, DE-A 10313214, DE-A 10313213, DE-A 10313212, DE-A 10308824, DE-A 10313208 and DE-A 10211275 is. The oxidizing agent may be pure molecular oxygen, Air, oxygenated air or any other mixture be added from oxygen and inert gas.

In the process according to the invention, the composition of the feed gas mixture (compare DE-A 10245585 and DE-A 10246119) for the reaction zone B is adjusted so that it satisfies the following molar ratios:
Propane: propene: N ₂ : O ₂ : H ₂ O: other = 0, 5 to 20: 1: 0.1 to 40: 0.1 to 10: 0 to 20: 0 to 1.

With Advantage be the aforementioned molar ratios according to the invention = 1 or 2 to 10: 1: 0.5 to 20: 0.5 to 5: 0.01 to 10: 0 to 1.

Cheap is it also, if the aforementioned molar ratios according to the invention = 3 bis 6: 1: 1 to 10: 1 to 3: 0.1 to 2: 0 to 0.5.

In in a known manner runs the heterogeneously catalyzed gas phase partial oxidation of propylene to acrylic acid with molecular oxygen in two along the reaction coordinate successive steps, of which the first to the acrolein and the second leads from acrolein to acrylic acid.

This reaction sequence in two successive steps opens in a conventional manner the possibility of the reaction zone B of the process according to the invention in this case two stages, ie, execute in two consecutively arranged oxidation zones, wherein in each of the two oxidation the oxidic catalyst to be used can be optimally adjusted (also the reaction can be stopped at the acrolein stage). Thus, for the first oxidation zone (propylene → acrolein), preference is generally given to using a catalyst based on the multimetal oxides containing the element combination Mo-Bi-Fe, while for the second oxidation zone (acrolein → acrylic acid) catalysts based on the element combination Mo -V containing multimetal oxides are preferred.

Appropriate Multimetal oxide catalysts for the two oxidation zones are often described above and the Expert well known. For example, EP-A 253 409 Page 5 for related US patents.

Cheap catalysts for the Both oxidation zones also disclose DE-A 4 431 957, DE-A 102004025445 and DE-A 4431949. This applies in particular to those of the general formula I in the two aforementioned publications. Particularly advantageous catalysts for the Both oxidation zones reveal the specifications DE-A 10325488, DE-A 10325487, DE-A 10353954, DE-A 10344149, DE-A 10351269, DE-A 10350812, DE-A 10350822.

The The abovementioned documents also disclose suitable ones according to the invention Gasphasenpartialoxidationsverfahren.

For the first Partial oxidation step, heterogeneously catalyzed gas-phase partial oxidation from propylene to acrolein, come, as already said, in principle all Mo, Bi and Fe containing Multimetalloxidmassen into consideration.

This In particular, the multimetal oxide active materials are the general ones Formula I of DE-A 19955176, the multimetal oxide active materials of the general Formula I of DE-A 19948523, the Multimetalloxidaktivmassen the general Formulas I, II and III of DE-A 10101695, the multimetal oxide active compounds of the general formulas I, II and III of DE-A 19948248 and the Multimetalloxidaktivmassen the general formulas I, II and III of DE-A 19955168 and the in EP-A 700714 mentioned Multimetalloxidaktivmassen.

Also suitable for this oxidation step are the multimetal oxide catalysts containing Mo, Bi and Fe which are described in DE-A 10046957, DE-A 10063162, DE-C 3338380, DE-A 19902562, EP-A 15565, DE-C 2380765, EP-A A 807465, EP-A 279374, DE-A 3300044, EP-A 575897, US-A 4438217, DE-A 19855913, WO 98/24746, DE-A 19746210 (those of the general formula II), JP-A 91 / 294239, EP-A 293224 and EP-A 700714 are disclosed. This applies in particular to the exemplary Ausfüh tion forms in these documents, among which those of EP-A 15565, EP-A 575897, DE-A 19746210 and DE-A 19855913 are particularly preferred. Particularly noteworthy in this context are a catalyst according to Example 1 c of EP-A 15565 and a catalyst to be prepared in a corresponding manner, the active composition but the composition Mo ₁₂ Ni _6.5 Zn ₂ Fe ₂ Bi ₁ P _0.0065 K _{0, 06} O _x · 10SiO ₂ . Further mention should be made of the example with the running No. 3 from DE-A 19855913 (stoichiometry: Mo ₁₂ Co ₇ Fe ₃ Bi _0.6 K _0.08 Si _1.6 O _x ) as a hollow cylinder full catalyst of the geometry 5 mm × 3 mm × 2 mm (outer diameter × height × inner diameter) and the multimetal oxide II - full catalyst according to Example 1 of DE-A 19746210. Furthermore, the multimetal oxide catalysts of US Pat. No. 4,438,217 may be mentioned. The latter applies in particular when these hollow cylinders have a geometry of 5.5 mm × 3 mm × 3.5 mm, or 5 mm × 2 mm × 2 mm, or 5 mm × 3 mm × 2 mm, or 6 mm × 3 mm × 3 mm, or 7 mm × 3 mm × 4 mm (outer diameter × height × inner diameter). Other possible catalyst geometries in this context are strands (eg 7.7 mm in length and 7 mm in diameter, or 6.4 mm in length and 5.7 mm in diameter).

A multiplicity of the multimetal oxide active compositions suitable for the step from propylene to acrolein can be obtained under the general formula IV Mo ₁₂ Bi _a Fe _b X ¹ _c X ² _d X ³ _e X ⁴ _f O _n (IV) in which the variables have the following meaning:
X ¹ = nickel and / or cobalt,
X ² = thallium, an alkali metal and / or an alkaline earth metal,
X ³ = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten,
X ⁴ = silicon, aluminum, titanium and / or zirconium,
a = 0.5 to 5,
b = 0.01 to 5, preferably 2 to 4,
c = 0 to 10, preferably 3 to 10,
d = 0 to 2, preferably 0.02 to 2,
e = 0 to 8, preferably 0 to 5,
f = 0 to 10 and
n = a number determined by the valence and frequency of the elements other than oxygen in IV,
subsumed.

she are available in a manner known per se (see, for example, DE-A 4023239) and are customary in substance to spheres, rings or cylinders shaped or in Shape of shell catalysts, i.e. coated with the active composition preformed, inert support bodies used. Of course can you but also be used in powder form as catalysts.

In principle, active compounds of the general formula IV can be prepared in a simple manner by producing an intimate, preferably finely divided, stoichiometrically composed, dry mixture of suitable sources of their elemental constituents and calcining this at temperatures of 350 to 650 ° 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) and also under a reducing atmosphere (eg., Mixture of inert gas, NH ₃ , CO and / or H ₂ ) take place. The calcination time can be several minutes to several hours and usually decreases with temperature. Suitable sources of the elemental constituents of the multimetal oxide active compositions IV 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, suitable starting compounds in particular are halides, nitrates, formates, oxalates, citrates, acetates, carbonates, amine complexes, ammonium salts and / or hydroxides (compounds such as NH ₄ OH, (NH ₄ ) ₂ CO ₃ , NH ₄ NO ₃ , NH ₄ CHO ₂ , CH ₃ OOOH, NH ₄ CH ₃ CO ₂ and / or ammonium oxalate, which can disintegrate and / or decompose into gaseous compounds at the latest during later calcination, can be additionally incorporated into the intimate dry mixture) ,

The intimate mixing of the starting compounds for the preparation of multimetal oxide active materials IV can be done in dry or wet form. It happens in dry Form, the starting compounds are expediently as finely divided Powder used and after mixing and optionally compacting subjected to calcination. Preferably, the intimate mixing takes place but in wet form. Usually be the starting compounds in the form of an aqueous solution and / or suspension mixed together. Especially intimate dry mixtures are then obtained in the described mixing method, if only from in dissolved Form present sources of elementary constituents becomes. As a solvent Water is preferably used. Subsequently, the resulting aqueous mass dried, the drying process preferably by spray drying the aqueous Mixture with outlet temperatures of 100 to 150 ° C takes place.

The Multimetal oxide active compounds of the general formula IV can be used for the step "propylene → acrolein" both in powder form used as well as molded to specific catalyst geometries The shaping before or after the final calcination can be done. For example, you can from the powder form of the active composition or its uncalcined and / or partially calcined precursor by compacting to the desired Catalyst geometry (eg by tabletting, extruding or Extrusion) Vollkatalysatoren be prepared, where appropriate Aids such. As graphite or stearic acid as a lubricant and / or Molding aids and reinforcing agents such as glass microfibers, asbestos, silicon carbide or potassium titanate can be added. Suitable Vollkatalysatorgeometrien are z. B. solid cylinder or hollow cylinder with an outer diameter and a length from 2 to 10 mm. In the case of the hollow cylinder is a wall thickness of 1 to 3 mm appropriate. Of course you can the full catalyst also have spherical geometry, wherein the ball diameter 2 to 10 mm can be.

A especially cheap Hollow cylinder geometry is 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter), especially in the case of full catalysts.

Of course, the shaping of the pulverulent active composition or its pulverulent, not yet and / or partially calcined, precursor composition can also be effected by application to preformed inert catalyst supports. The coating of the carrier body for the preparation of the coated catalysts is usually carried out in a suitable rotatable container, as z. B. from DE-A 2909671, EP-A 293859 or from EP-A 714700 is known. Appropriately, the carrier body is used for coating moistened the applied powder mass and after application, for. B. by means of hot air, dried again. The layer thickness of the powder mass applied to the carrier body is expediently chosen to be in the range 10 to 1000 μm, preferably in the range 50 to 500 μm and particularly preferably in the range 150 to 250 μm.

When support materials can usual porous or nonporous Aluminas, silica, thoria, zirconia, silicon carbide or silicates such as magnesium or aluminum silicate are used. They are behaving the method of the invention underlying target reaction is generally substantially inert. The carrier bodies can be regular or irregular shaped being regularly shaped Carrier body with clearly formed surface roughness, z. As balls or hollow cylinders, are preferred. Suitable is the Use of substantially non-porous, surface rough, spherical carriers Steatite whose diameter is 1 to 10 mm or 8 mm, preferably 4 to 5 mm. But is also suitable for the use of cylinders as a carrier body whose Length 2 up to 10 mm and their outer diameter 4 to 10 mm. In the case of fiction, according to suitable Rings as a carrier body lies the wall thickness over it usually at 1 to 4 mm. According to the invention preferred to be used annular Own carrier body a length from 2 to 6 mm, an outer diameter from 4 to 8 mm and a wall thickness of 1 to 2 mm. Suitable according to the invention are above all rings of geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) as a carrier body. The Fineness of the surface of the carrier body to be applied catalytically active oxide masses will of course be of the desired shell thickness adapted (see EP-A 714 700).

For the step from propylene to acrolein to be used Multimetalloxidaktivmassen are also compounds of the general formula V, [Y ¹ _{a '} Y ² _b' O _{x '} ] _p [Y ³ _c' Y ⁴ _{d '} Y ⁵ _e' Y ⁶ _f Y ⁷ _{g '} Y ² _h' O _{y '} ] _q (V) in which the variables have the following meaning:
Y ¹ = only bismuth or bismuth and at least one of the elements tellurium, antimony, tin and copper,
Y ² = molybdenum or molybdenum and tungsten,
Y ³ = an alkali metal, thallium and / or samarium,
Y ⁴ = an alkaline earth metal, nickel, cobalt, copper, manganese, zinc, tin, cadmium and / or mercury,
Y ⁵ = iron or iron and at least one of the elements chromium and cerium,
Y ⁶ = phosphorus, arsenic, boron and / or antimony,
Y ⁷ = a rare earth metal, titanium, zirconium, niobium, tantalum, rhenium, ruthenium, rhodium, silver, gold, aluminum, gallium, indium, silicon, germanium, lead, thorium and / or uranium,
a '= 0.01 to 8,
b '= 0.1 to 30,
c '= 0 to 4,
d '= 0 to 20,
e '> 0 to 20,
f '= 0 to 6,
g '= 0 to 15,
h '= 8 to 16,
x ', y' = numbers which are determined by the valence and frequency of the elements other than oxygen in V, and
p, q = numbers whose ratio p / q is 0.1 to 10, containing three-dimensionally expanded regions of the chemical composition Y ¹ _{a '} Y ² _b' O _x, demarcated from their local environment due to their different composition from their local environment _' whose largest diameter (longest through the center of gravity of the area going direct link of two located on the surface (interface) of the area points) 1 nm to 100 microns, often 10 nm to 500 nm or 1 micron to 50 or 25 microns.

Particularly advantageous multimetal oxide compositions V are those in which Y ¹ is only bismuth.

Among these, those which are of the general formula VI, [Bi _{a ''} Z ² _{b ''} O _{x ''} ] _{p ''} [Z ² ₁₂ Z ³ _{c ''} Z ⁴ _{d ''} Fe _{e ''} Z ⁵ _{f ''} Z ⁶ _{g ''} Z ⁷ _{h ''} O _y'' ] _q'' (VI), in which the variants have the following meaning:
Z ² = molybdenum or molybdenum and tungsten,
Z ³ = nickel and / or cobalt,
Z ⁴ = thallium, an alkali metal and / or an alkaline earth metal,
Z ⁵ = phosphorus, arsenic, boron, antimony, tin, cerium and / or lead,
Z ⁶ = silicon, aluminum, titanium and / or zirconium,
Z ⁷ = copper, silver and / or gold,

a '' = 0.1 to 1,
b '' = 0.2 to 2,
c '' = 3 to 10,
d "= 0.02 to 2,
e "= 0.01 to 5, preferably 0.1 to 3,
f '' = 0 to 5,
g '' = 0 to 10,
h '' = 0 to 1,
x '', y '' = numbers determined by the valency and frequency of the element other than oxygen in VI,
p '', q '' = numbers whose ratio p '' / q '' is 0.1 to 5, preferably 0.5 to 2, wherein those compounds VI are very particularly preferred in which
Z ² _{b ''} = (tungsten) _{b ''} and Z ² ₁₂ = (molybdenum) ₁₂ .

Furthermore, it is advantageous if at least 25 mol% (preferably at least 50 mol% and particularly preferably 100 mol%) of the total fraction [Y ¹ _{a '} Y ² _b' O _{x '} ] _p ([Bi _a'' Z ² _{b "} O _x" ] _{p "} ) of the multimetal oxide materials V (multimetal oxide materials VI) which are suitable according to the invention in the multimetal oxide compositions V (multimetal oxide compositions VI) which are suitable according to the invention in the form of three-dimensionally extended chemical substances which differ from their local environment because of their local environment Composition of demarcated, areas of chemical composition
Y ¹ _{a '} Y ² _b' O _{x '} [Bi _a'' Z ² _b'' O _x'' ] are present whose largest diameter in the range 1 nm to 100 microns.

Regarding the shaping applies with respect Multimetal oxide V-catalysts in the multimetal oxide materials IV catalysts said.

The Preparation of Multimetal Oxide Compounds V-Active Compounds is, for. In EP-A 575897 and in DE-A 19855913 described.

The The above-recommended inert support materials include i.a. also as inert materials for dilution and / or delimitation of corresponding fixed catalyst beds, or as their protective and / or the gas mixture heating pre-feed into consideration.

At this point should be mentioned that yourself all catalysts and multimetal oxide materials used for the step from propylene to acrolein were also recommended in principle for the Partialammoxidation of propylene to acrylonitrile are suitable.

For the second Step, the heterogeneously catalyzed gas-phase partial oxidation of Acrolein to acrylic acid, come, as already said, in principle all Mo and V containing Multimetal oxide compositions as active compounds into consideration, for. B. those of DE-A 10046928.

A variety of the same, for. B. those of DE-A 19815281, can be found under the general formula VII, Mo ₁₂ V _a X ¹ _b X ² _c X ³ _d X ⁴ _e X ⁵ _f X ⁶ _g O _n (VII) in which the variables have the following meaning:
X ¹ = W, Nb, Ta, Cr and / or Ce,
X ² = Cu, Ni, Co, Fe, Mn and / or Zn,
X ³ = Sb and / or Bi,
X ⁴ = one or more alkali metals,
X ⁵ = one or more alkaline earth metals,
X ⁶ = Si, Al, Ti and / or Zr,
a = 1 to 6,
b = 0.2 to 4,
c = 0.5 to 18,
d = 0 to 40,
e = 0 to 2,
f = 0 to 4,
g = 0 to 40 and
n = a number which is determined by the valence and frequency of the elements other than oxygen in VII,
subsumed.

Preferred embodiments within the active multimetal oxides VII according to the invention are those which are covered by the following meanings of the variables of general formula VII:
X ¹ = W, Nb, and / or Cr,
X ² = Cu, Ni, Co, and / or Fe,
X ³ = Sb,
X ⁴ = Na and / or K,
X ⁵ = Ca, Sr and / or Ba,
X ⁶ = Si, Al, and / or Ti,
a = 1.5 to 5,
b = 0.5 to 2,
c = 0.5 to 3,
d = 0 to 2,
e = 0 to 0.2,
f = 0 to 1 and
n = a number determined by the valence and frequency of the elements other than oxygen in VII.

However, very particularly preferred multimetal oxides VII according to the invention are those of the general formula VIII Mo ₁₂ V _{a '} Y ¹ _b' Y ² _{c '} Y ⁵ _f Y ⁶ _g' O _{n '} (VIII), With
Y ¹ = W and / or Nb,
Y ² = Cu and / or Ni,
Y ⁵ = Ca and / or Sr,
Y ⁶ = Si and / or Al,
a '= 2 to 4,
b '= 1 to 1.5,
c '= 1 to 3,
f = 0 to 0.5
g '= 0 to 8 and
n '= a number which is determined by the valence and frequency of the elements other than oxygen in VIII.

The suitable according to the invention Multimetal oxide active materials (VII) are known per se, for. B. available in DE-A 4335973 or in EP-A 714700.

In principle, suitable for the step "acrolein → acrylic acid" suitable Multimetalloxidaktivmassen, especially those of the general formula VII, in a simple manner be prepared by suitable sources of their elemental constituents an intimate, preferably finely divided, their stoichiometry correspondingly composed, dry mixture and this calcined at temperatures of 350 to 600 ° 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 a reducing atmosphere (eg., Mixtures of inert gas and reducing gases such as H ₂ , NH ₃ , CO, methane and / or acrolein or said reducing acting gases per se) performed become. The calcination time can be several minutes to several hours and usually decreases with temperature. Suitable sources of the elemental constituents of the multimetal oxide active compounds VII 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.

The intimate mixing of the starting compounds for the preparation of multimetal oxide materials VII can be done in dry or wet form. Is it done in dry form, so the starting compounds are suitably used as finely divided powder and after mixing and optionally Compacting subjected to calcination. Preferably, this is done intimate mixing but in wet form.

Usually be the starting compounds in the form of an aqueous solution and / or suspension mixed together. Especially intimate dry mixtures are then obtained in the described mixing method, if only from in dissolved Form present sources of elementary constituents becomes. As a solvent Water is preferably used. Subsequently, the resulting aqueous mass dried, the drying process preferably by spray drying the aqueous Mixture with outlet temperatures of 100 to 150 ° C takes place.

The resulting Multimetalloxidmassen, in particular those of the general Formula VII, can for the Acrolein oxidation both in powder form and to certain catalyst geometries be used, wherein the shaping before or after the final Calcination can be done. For example, from the powder form of Active mass or its uncalcined precursor mass by compaction to the desired Catalyst geometry (eg by tabletting, extruding or Extrusion) Vollkatalysatoren be prepared, where appropriate Aids such. As graphite or stearic acid as a lubricant and / or Molding aids and reinforcing agents such as glass microfibers, asbestos, silicon carbide or potassium titanate can be added. Suitable Vollkatalysatorgeometrien are z. B. solid cylinder or Hollow cylinder with an outer diameter and a length from 2 to 10 mm. In the case of the hollow cylinder is a wall thickness of 1 to 3 mm appropriate. Of course you can the full catalyst also have spherical geometry, wherein the ball diameter 2 to 10 mm (eg 8.2 mm or 5.1 mm).

Of course you can the shape of the powdered Active mass or its powdery, yet uncalcined, precursor material also be carried out by applying to preformed inert catalyst support. The coating of the carrier body for Preparation of the coated catalysts is usually in a suitable rotatable container executed as it is z. Example from DE-A 2909671, EP-A 293859 or from EP-A 714700 is known.

Conveniently, is used to coat the carrier body applied powder compound moistened and after application, z. B. by means of hotter Air, dried again. The layer thickness of the applied to the carrier body Powder mass is expediently in the range 10 to 1000 μm, preferably in the range 50 to 500 microns and more preferably in the range 150 to 250 microns, selected.

When support materials can usual porous or nonporous Aluminas, silica, thoria, zirconia, silicon carbide or silicates such as magnesium or aluminum silicate are used. The carrier bodies can be regular or irregular shaped be regularly formed carrier body with clearly formed surface roughness, z. As balls or hollow cylinder with chippings, are preferred. Suitable is the use of essentially nonporous, surface roughening, spherical carriers made of steatite whose diameter is 1 to 10 mm or 8 mm, preferably 4 to 5 mm. that is, suitable ball geometries can Diameter of 8.2 mm or 5.1 mm. But it is suitable also the use of cylinders as carrier bodies whose length is 2 to 10 mm and its outer diameter 4 to 10 mm. In the case of rings as a carrier body is the wall thickness over it usually at 1 to 4 mm. Preferably to be used annular carrier body have a length of 2 to 6 mm, an outer diameter from 4 to 8 mm and a wall thickness of 1 to 2 mm. Are suitable especially rings of geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) as a carrier body. The Fineness of the surface of the carrier body to be applied catalytically active oxide masses will of course be of the desired shell thickness adapted (see EP-A 714 700).

Favorable for the step "acrolein → acrylic acid" to be used Multimetalloxidaktivmassen are also compounds of general formula IX, [D] _P [E] _q (IX), in which the variables have the following meaning:
D = Mo ₁₂ V _{a "} Z ¹ _b" Z ² _{c "} Z ³ _d" Z ⁴ _{e "} Z ⁵ _f" Z ⁶ _{g "} O _x" ,
E = Z ⁷ ₁₂ Cu _{h "} H _i" O _{y "} ,
Z ¹ = W, Nb, Ta, Cr and / or Ce,
Z ² = Cu, Ni, Co, Fe, Mn and / or Zn,
Z ³ = Sb and / or Bi,
Z ⁴ = Li, Na, K, Rb, Cs and / or H
Z ⁵ = Mg, Ca, Sr and / or Ba,
Z ⁶ = Si, Al, Ti and / or Zr,
Z ⁷ = Mo, W, V, Nb and / or Ta, preferably Mo and / or W.
a '' = 1 to 8,
b '' = 0.2 to 5,
c '' = 0 to 23,
d '' = 0 to 50,
e '' = 0 to 2,
f '' = 0 to 5,
g '' = 0 to 50,
h '' = 4 to 30,
i '' = 0 to 20 and
× '', y '' = numbers which are determined by the valence and frequency of the element different from oxygen in IX and
p, q = non-zero numbers whose ratio p / q is 160: 1 to 1: 1 and which are obtainable by using a multimetal oxide composition E Z ⁷ ₁₂ Cu _{h "} H _i" O _{y "} (E), in finely divided form vorbildet (starting material 1) and then the preformed solid starting material 1 in an aqueous solution, an aqueous suspension or in a finely divided dry mixture of sources of elements Mo, V, Z ¹ , Z ² , Z ³ , Z ⁴ , Z ⁵ , Z ⁶ , which the aforementioned elements in the stoichiometry D MO ₁₂ V _{a "} Z ¹ _b" Z ² _{c "} Z ³ _d" Z ⁴ _{e "} Z ⁵ _f" Z ⁶ _{g "} (D) contains (starting material 2), in the desired ratio p: q incorporated, which optionally dried resulting aqueous mixture, and the dry precursor material thus calcined prior to or after drying to the desired catalyst geometry at temperatures of 250 to 600 ° C.

Prefers are the Multimetalloxidmassen IX, in which the incorporation of Preformed solid starting material 1 in an aqueous starting material 2 at a temperature <70 ° C is carried out. A detailed description of the preparation of multimetal oxide compositions VI catalysts contain z. For example, EP-A 668104, DE-A 19736105, DE-A 10046928, DE-A 19740493 and DE-A 19528646.

Regarding the shaping applies with respect Multimetal oxide materials IX catalysts that in the case of the multimetal oxide materials VII catalysts said.

For the step "acrolein → acrylic acid" in excellent Suitable multimetal oxide catalysts are also those of DE-A 19815281, in particular with multimetal oxide active compounds of the general Formula I of this document.

Advantageous be for the step from propylene to acrolein full-catalyst rings and for the step from Acrolein to acrylic acid Shell catalyst rings used.

The execution of the first step of the partial oxidation, from propylene to acrolein, can with the described catalysts z. B. in a one-zone Vielkontaktrohr fixed bed reactor take place as described in DE-A 4431957.

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

Frequently becomes for a propylene: oxygen: inert gases (including water vapor) Volume (NI) ratio of 1: (1.0 to 3.0) :( 5 to 25), preferably 1: (1.7 to 2.3) :( 10 to 15) worked. The reaction pressure is usually in the range of 1 to 3 bar and the total space load is preferably 1500 to 4000 or 6000 Nl / l · h or more. The propylene load is typically 90 to 200 Nl / l · h or up to 300 Nl / l · h or more.

Preferably, the single-zone multi-contact fixed-bed reactor with the feed gas mixture of flowed on top. As a heat exchange agent is advantageously a molten salt, preferably consisting of 60 wt .-% potassium nitrate (KNO ₃ ) and 40 wt .-% sodium nitrite (NaNO ₂ ), or from 53 wt .-% potassium nitrate (KNO ₃ ), 40 wt .-% Sodium nitrite (NaNO ₂ ) and 7 wt .-% sodium nitrate (NaNO ₃ ) used.

On the Reactor can be considered Salt melt and reaction gas mixture both in the DC and conducted in countercurrent become. The molten salt itself is preferably guided meander-shaped around the catalyst tubes.

• – zunächst auf einer Länge von 40 bis 80 bzw. bis 60 % der Kontaktrohrlänge entweder nur Katalysator, oder ein Gemisch aus Katalysator und Inertmaterial, wobei letzteres, bezogen auf die Mischung, einen Gewichtsanteil von bis zu 20 Gew.-% ausmacht (Abschnitt C);
• – daran anschließend auf einer Länge von 20 bis 50 bzw. bis 40 % der Gesamtrohrlänge entweder nur Katalysator, oder ein Gemisch aus Katalysator und Inertmaterial, wobei letzteres, bezogen auf die Mischung, einen Gewichtsanteil von bis zu 40 Gew.-% ausmacht (Abschnitt B); und
• – abschließend auf einer Länge von 10 bis 20 % der Gesamtrohrlänge eine Schüttung aus Inertmaterial (Abschnitt A), die vorzugsweise so gewählt wird, dass sie einen möglichst geringen Druckverlust bedingt.

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

• - First, on a length of 40 to 80 or 60% of the catalyst tube length either catalyst, or a mixture of catalyst and inert material, the latter, based on the mixture, a weight fraction of up to 20 wt .-% makes (Section C );
• - Thereafter, over a length of 20 to 50 or 40% of the total tube length, either catalyst, or a mixture of catalyst and inert material, the latter, based on the mixture, a weight fraction of up to 40 wt .-% makes (Section B); and
• - Finally, over a length of 10 to 20% of the total tube length, a bed of inert material (Section A), which is preferably chosen so that it causes the lowest possible pressure loss.

Prefers section C is undiluted.

The aforementioned Loading variant is particularly useful if as catalysts such example 1 of DE-A 10046957 or according to Example 3 of DE-A 10046957 and as an inert material rings of steatite with the geometry 7 mm × 7 mm × 4 mm (outer diameter × height × inner diameter) used become. With regard to the salt bath temperature, the same applies in DE-A 4431957 said.

The execution of the first step of the partial oxidation, from propylene to acrolein, can with the described catalysts but also z. In one Two-zone multiple catalyst tube fixed bed reactor take place, as described in DE-A 19910506. In both above described cases the achieved propene conversion is usually in a single pass at values ≥ 90 mol%, or ≥95 mol%. The implementation of the second Step of the partial oxidation, from acrolein to acrylic acid, can with the described catalysts z. B. in a one-zone Vielkontaktrohr fixed bed reactor carried out become, as it DE-A 4431949 describes. In general, that will Product mixture of Propylenpartialoxidation to acrolein as such (optionally after the intermediate cooling thereof), that is, without Nebenkomponentenabtrennung, led in the acrolein oxidation to acrylic acid.

Of the for the second step of the partial oxidation required oxygen is thereby preferably added as air and usually the product gas mixture the Propylenpartialoxidation immediately added (but he can even in the starting reaction gas mixture for the Propylenpartialoxidation be included).

In The rule is the feed gas mixture of such Acroleinpartialoxidation then composed as follows: acrolein: oxygen: water vapor : Inert gas volume ratio (Nl) of 1: (1 to 3) :( 0 to 20) :( 3 to 30), preferably 1: (1 to 30) 3) :( 0.5 to 10) :( 7 to 18).

Of the Reaction pressure is usually also here 1 to 3 bar and the total space load is preferably at 1000 to 3800, or 4800 Nl / l · h or more. The acrolein load is typically 80 to 190, or 290 Nl / l · h or more.

Preferably, the single-zone multiple contact tube fixed bed reactor is also supplied with the feed gas mixture from above. As a heat exchange agent, a salt melt, preferably from 60% by weight of potassium nitrate (KNO ₃ ) and 40% by weight of sodium nitrite (NaNO ₂ ) or of 53% by weight of potassium nitrate (KNO ₃ ) is expediently also used in the second stage. 40 wt .-% sodium nitrite (NaNO ₂ ) and 7 wt .-% sodium nitrate (Na-NO ₃ ) consisting used. Viewed over the reactor molten salt and reaction gas mixture can be performed both in DC and in countercurrent. The molten salt itself is preferably guided meander-shaped around the catalyst tubes.

• – zunächst auf einer Länge von 50 bis 80 bzw. bis 70 % der Kontaktrohrlänge entweder nur Katalysator oder ein Gemisch aus Katalysator und Inertmaterial, wobei letzteres, bezogen auf die Mischung, einen Gewichtsanteil von bis zu 20 Gew.-% ausmacht (Abschnitt C);
• – daran anschließend auf einer Länge von 20 bis 40 % der Gesamtrohrlänge entweder nur Katalysator, oder ein Gemisch aus Katalysator und Inertmaterial, wobei letzteres, bezogen auf die Mischung, einen Gewichtsanteil von bis zu 50 bzw. bis zu 40 Gew.-% ausmacht (Abschnitt B); und
• – abschließend auf einer Länge von 5 bis 20 % der Gesamtrohrlänge eine Schüttung aus Inertmaterial (Abschnitt A), die vorzugsweise so gewählt wird, dass sie einen möglichst geringen Druckverlust bedingt.

If the contact tubes are flown from top to bottom, it is advisable to feed the contact tubes from bottom to top as follows:

• - First on a length of 50 to 80 or up to 70% of the contact tube length either catalyst only or a mixture of catalyst and inert material, the latter, based on the mixture, a weight fraction of up to 20 wt .-% makes (Section C);
• - Thereafter, over a length of 20 to 40% of the total tube length, either catalyst only, or a mixture of catalyst and inert material, the latter, based on the mixture, a weight fraction of up to 50 and up to 40 wt .-% makes ( Section B); and
• - Finally, on a length of 5 to 20% of the total tube length, a bed of inert material (Section A), which is preferably chosen so that it causes the lowest possible pressure loss.

Prefers section C is undiluted. As is generally the case with heterogeneously catalyzed gas phase partial oxidation from acrolein to acrylic acid (especially at high acrolein loadings of the catalyst bed and high water vapor contents of the feed gas mixture), can Section B also consists of two consecutive catalyst dilutions exist (for the purpose of minimizing hotspot temperature and hotspot temperature sensitivity). From bottom to top first with up to 20 wt .-% inert material and subsequently with> 20 wt .-% to 50 or to 40 wt .-% inert material. The section C is then preferred undiluted.

For a flow of the Contact tubes from bottom to top, is the contact tube feed in expediently vice versa.

The aforementioned Loading variant is particularly useful if as catalysts such according to the preparation example 5 of DE-A 10046928 or those according to DE-A 19815281 and as inert material Steatite rings with the geometry 7 mm × 7 mm × 4 mm or 7 mm × 7 mm × 3 mm (each Outside diameter × height × inside diameter) be used. With regard to the salt bath temperature, this applies in the DE-A 44 319 49. It is usually chosen so that the scored Acrolein conversion in a single pass normally ≥ 90 mol%, or ≥ 95 mol% or 99 mol%.

The execution of the second step of the partial oxidation, from acrolein to acrylic acid with the described catalysts but also z. B. in a two-zone Vielkontaktrohr fixed bed reactor take place, as described in DE-19910508. For acrolein sales this applies mentioned above. Also in the case of carrying out this second step in a two-zone multi-contact fixed bed reactor, one becomes Production of the feed gas mixture expedient directly the product gas mixture a directed to the first step partial oxidation (optionally after intercooling the same) (as described above). Of the for the second step of the partial oxidation required oxygen is thereby preferably added as air and in the second case the product gas mixture the first step of the partial oxidation is added immediately.

at a two-stage driving style with immediate further use the product gas mixture of the first step of the partial oxidation to feed the second step of the partial oxidation usually two single-zone Vielkontaktrohr fixed bed reactors (at high Reaktandenbelastung the catalyst bed is, as generally valid, a DC mode between reaction gas and salt bath (heat transfer medium) via the tube bundle reactor considered preferred) of the two two-zone Vielkontaktrohr fixed bed reactors in series. A mixed series connection (one-zone / two-zone or vice versa) is also possible.

Between The reactors may be an intercooler, if necessary Inert fillings included can, which can exercise a filter function. The salt bath temperature of multi-contact tube reactors for the first step of the partial oxidation of propylene to acrylic acid is in the Usually 300 to 400 ° C. The salt bath temperature of multi-contact tube reactors for the second Step of the partial oxidation of propylene to acrylic acid, the Partial oxidation of acrolein to acrylic acid, is usually 200 to 350 ° C. In addition, the Heat exchange means (preferably molten salts) normally in such amounts the relevant multi-contact fixed bed reactors that led the difference between their input and their output temperature usually ≤ 5 ° C. As already mentioned, can however, both steps of the partial oxidation of propylene to acrylic acid also as described in DE-A 10121592 in a reactor on a Charging be done.

It be mentioned again, the existence Part of the feed gas mixture for the first step ("propylene → acrolein") from the partial oxidation may be incoming cycle gas (residual gas).

This is, as already said, a molecular oxygen-containing gas, which after the target product separation (acrolein and / or acrylic acid separation) from the product gas mixture of Partialoxi dation and can be partially recycled as an inert diluent gas in the feed for the first and / or optionally second step of the partial oxidation of propylene to acrolein and / or acrylic acid.

Prefers Will such propane, molecular oxygen and possibly not reacted propylene-containing cycle gas (residual gas) but exclusively on way according to the invention recycled to the reaction zone A of the process according to the invention.

It may also be mentioned that is a partial oxidation according to the invention so performed that can be done first a reaction gas mixture over directs the catalyst feed containing no oxygen. In this Case will be for required the partial oxidation Oxygen is provided as lattice oxygen. In a following Regeneration step with an oxygen-containing gas (eg air, with Oxygen enriched air or deoxygenated air) the catalyst bed is regenerated to subsequently turn on for a Oxygen-free reaction gas mixture to be available, etc.

In summary again forms a tube bundle reactor within which the catalyst feed along the individual contact tubes with completion of the first reaction step changes accordingly (Such as reaction zone B according to the invention suitable Propylenpartialoxidationen teach z. For example, EP-A 911313, EP-A 979813, EP-A 990636 and DE-A 2830765) the simplest realization form of two oxidation zones for the both steps of the partial oxidation of propylene to acrylic acid. Possibly In this case, the feed of the catalyst tubes with catalyst an inert bed interrupted.

Preferably However, the two oxidation zones are in the form of two consecutively switched tube bundle systems realized. these can are in a reactor, the transition from a tube bundle to other tube bundle from a not in the contact tube accommodated (appropriately walkable) fill be formed from inert material. While the contact tubes in the Usually be lapped by a heat transfer medium, this does not reach an inert bed attached as described above. Advantageously, therefore, the two catalyst tube bundles are spatially separated from each other Reactors housed. As a rule, this is between the two tube bundle reactors an intercooler, an optional acrolein afterburning in the product gas mixture, which leaves the first oxidation zone to reduce. Instead of Tube reactors can also plate heat exchanger reactors with salt and / or boiling cooling, as they are z. For example, DE-A 19 929 487 and DE-A 19 952 964 describe be used.

The Reaction temperature in the first oxidation zone is usually at 300 to 450 ° C, preferably at 320 to 390 ° C. The reaction temperature in the second oxidation zone is in the Usually at 200 to 370 ° C, often at 220 to 330 ° C. The reaction pressure is in both oxidation zones appropriate 0.5 to 5, advantageously 1 to 3 bar. The load (Nl / l · h) of Oxidation catalysts with reaction gas is in both oxidation zones often 1500 to 2500 h-1 or up to 4000 h-1. The load with popylene can are at 100 to 200 or 300 and more Nl / l · h.

in principle can the two oxidation zones designed in the method according to the invention be like z. In DE-A 19 837 517, DE-A 19 910 506, DE-A 19 910 508 and DE-A 19 837 519 is described. Usually, the external Tempering in the two oxidation zones, optionally in multi-zone reactor systems, in a conventional manner to the specific reaction gas mixture composition and catalyst feed adjusted.

Of the for the required according to the invention Reaction zone B as a total oxidant required molecular Oxygen may be in the case of a two-stage propylene partial oxidation to acrylic acid the feed gas mixture of the propylene oxidation in his Total amount to be added in advance. Of course, but also after the propylene oxidation are supplemented with oxygen. The latter is included the acrylic acid production prefers.

Preferably is in the first oxidation zone (propylene → acrolein) a molar ratio of propylene : molecular oxygen from 1: 1 to 3, often 1: 1.5 to 2 set. Similar numerical values are suitable for the molar ratio Acrolein: molecular oxygen in the second oxidation zone (1 : 0.5 to 1.5 would be preferred) for the partial oxidation of acrolein to acrylic acid.

In both oxidation zones, an excess of molecular oxygen generally has an advantageous effect on the kinetics of the gas phase oxidation. In contrast to the ratios in the reaction Zone A of the process according to the invention, the thermodynamic conditions in reaction zone B are essentially unaffected by the molar ratio of reactants, since the heterogeneously catalyzed gas phase partial oxidation of propylene to acrylic acid is subject to kinetic control. In principle, therefore, z. B. in the first oxidation zone and the propylene are presented to the molecular oxygen in molar excess. In this case, the excess propylene actually assumes the role of a diluent gas.

in principle is the heterogeneously catalyzed gas-phase partial oxidation of propylene to acrylic acid but also in a single oxidation zone feasible. In this In case of both reaction steps take place in an oxidation reactor which is charged with a catalyst, the implementation of both Catalysts can catalyze reaction steps. Of course it can be also the catalyst charge within the oxidation zone along the Change reaction coordinate continuously or abruptly. Of course, in one embodiment of the to be used according to the invention at least one partial oxidation in the form of two series-connected Oxidation zones from that leaving the first oxidation zone Product gas mixture contained selbigem in the first oxidation zone by-product, carbon monoxide and water vapor as needed before passing to the second oxidation zone partially or Completely be separated. Preferably, according to the invention, a procedure choose, which does not need such a separation.

As a source for the molecular oxygen required in the at least one partial oxidation, the z. B is added to the product gas mixture A 'prior to its use for charging the reaction zone B, both pure molecular oxygen and with inert gas such as CO ₂ , CO, noble gases, N ₂ and / or saturated hydrocarbons dilute molecular oxygen or nitric oxide into consideration.

In expedient manner will be at least to cover a partial requirement of molecular Use oxygen as the oxygen source.

By Metering of cold air to hot product gas mixture A 'may, in the context of inventive method also on direct way a cooling off be effected of the same.

in the Long-term operation of both reaction zones A, B will normally be a Reduction of the catalyst quality (especially activity and selectivity) associated with the catalysts used in both reaction zones. On the one hand, this can be counteracted by doing so each catalyst bed regenerated from time to time (this can in the oxidation zones z. B. also residual gas at elevated temperature over the respective catalyst bed of the oxidation catalysts are performed), as it is z. In DE-A 10351269, DE-A 10350812 and DE-A 10350822 is described. Also, the reaction temperature over the Operating time increased to compensate for a decreased catalyst activity. In surprising However, it has in both reaction zones A, B as advantageous proven, the respective working pressure in the gas phase, based on an identical load of the respective catalyst bed with reaction gas mixture in Nl / lh, while increase the service life of the catalyst bed to deactivate the Further counteract catalyst beds, as z. B. in DE-A 102004025445 is described. For this purpose, for example behind each of the applicable in the process according to the invention Reaction and separation zones a pressure regulator (in the simplest Case a throttle device, z. B. a throttle valve or a spin regulator) are attached (eg also partially permeable pinhole diaphragms, their holes successively partial or complete can be closed. Of course but can also only behind one of the zones such a Pressure control device are located. That is, in principle it is quite sufficient the pressure control device somewhere in the other flow path of the reaction gas, from where the pressure increase by backwater can propagate into the reaction zones. Typical aforementioned pressure increases (the continuously according to success Deactivation, but also can be done discontinuously) can Course of operation up to 3000 mbar and more.

In the case of a preparation of acrolein and / or acrylic acid, the product gas mixture B leaving the reaction zone B according to the invention is generally composed essentially of the target product acrolein or acrylic acid or its mixture with acrolein, unreacted molecular oxygen, propane, unreacted propylene, molecular Nitrogen, water vapor by-produced and / or as diluent gas, as by-product and / or diluent carbon oxides, as well as small amounts of other lower aldehydes, lower alkanecarboxylic acids (eg acetic acid, formic acid and propionic acid) and maleic anhydride, benzaldehyde, aromatic carboxylic acids and aromatic carboxylic acid anhydrides (e.g., phthalic anhydride and benzoate acid), optionally further hydrocarbons, such as. C4 hydrocarbons (eg, butene-1 and possibly other butenes) and other inert diluent gases.

The target product can be separated from the product gas mixture B in a manner known per se in a separation zone B (for example by partial or complete and optionally fractional condensation of the acrylic acid or by absorption of acrylic acid in water or in a high-boiling hydrophobic organic solvent or by absorption of acrolein in water or in aqueous solutions of lower carboxylic acids and subsequent workup of the condensates and / or absorbates; according to the invention, the product gas mixture B is preferably fractionally condensed, cf., for example, EP-A 1388533, EP-A 1388532, DE-A 10235847 EP-A 792867, WO 98/01415, EP-A 1015411, EP-A 1015410, WO 99/50219, WO 00/53560, WO 02/09839, DE-A 10235847, WO 03/041833, DE-A 10223058 DE-A 10243625, DE-A 10336386, EP-A 854129, US-A 4317926, DE-A 19837520, DE-A 19606877, DE-A 190501325, DE-A 10247240, DE-A 19740253, EP-A 695736 , EP-A 982,287, EP-A 1041062, EP-A 117146, DE-A 4308087, DE-A 4335172, DE-A 4436243, DE-A 19 924 532, DE-A 10332758 and DE-A 19 924 533). An acrylic acid removal can also be carried out as described in EP-A 982287, EP-A 982,289, DE-A 10336386, DE-A 10115277, DE-A 19606877, DE-A 19740252, DE-A 19627847, DE-A. A 10053086 and EP-A 982288 are made. Preferably, as in 7 WO / 0196271 or separated in the embodiments of the present application.

common Feature of the aforementioned separation method is (as already mentioned mentioned), that at the top of the respective separating internals containing Separation column, in the lower part of the product mixture B, usually after prior direct and / or indirect cooling thereof, normally a residual gas stream (main residual gas) remains, essentially those Components of the product gas mixture B contains, whose boiling point at atmospheric pressure (1 bar) ≤ -30 ° C (i.e. the hardly condensable or volatile constituents).

in the lower part of the separation column usually fall the heavier volatile Constituents of the product gas mixture B, including the respective target product, in condensed phase.

The Residual gas components are primarily propane, if necessary in the reaction zone B unreacted propylene, molecular oxygen as well as often the inert diluent gases used in the reaction zones A, B preferably in particular nitrogen and carbon dioxide. Steam can depending on the applied separation process in the residual gas only in traces or in amounts of up to 20% by volume or more.

From This (main) residual gas is inventively at least one (preferably (Main) residual gas composition) propane, molecular oxygen and optionally unreacted propylene (in reaction zone B) containing subset (preferably the total amount, optionally but only half, or two thirds, or three quarters of that total) as one Propane-containing feed stream is recycled to the reaction zone A. (Main) residual gas subsets can but also, as already described, recycled to the reaction zone B and / or burned for the purpose of generating energy.

at the work-up of the condensed phase (for the purpose of separation of the target product) can according to the invention (as already mentioned) other residual gases are incurred, as is usually attempted the total contained in the product gas mixture A amount of unreacted Propane in the reaction zone A and as part of the target separation regain. These usually still contain propane and optionally propylene, often but no more molecular oxygen (they still contain molecular Oxygen they are referred to in this document as minor residual gases). You can therefore optionally with the main residual gas to a total residual gas united in the reaction zone according to the invention A be returned. But it is also a separate utilization of such other remaining Gases possible. For example, you can as far as they are free of molecular oxygen, as a rule after compression, also as part of the reaction gas starting mixture the reaction zone A are recycled to the reaction zone A. Also, her Return in the reaction zone A take place elsewhere, if they are at molecular Oxygen are substantially free.

By the preferably complete Return of the (Main and / or secondary or total) residual gas from the separation zone B, which is an interconnection of several separation columns and optionally may comprise further separation apparatus different from columns, in the reaction zone A, takes place in continuous execution of inventive method such a continuous conversion of propane to acrylic acid and / or Acrolein.

Substantially according to the invention is that by the inventive repatriation of the (main and / or Secondary or total) residual gas from the separation zone B in the reaction zone A in the latter a transfer of Propane can be reached to propylene with almost 100% selectivity is.

The Advantageousness of the procedure according to the invention is thereby both at lower (≤ 30 mol%) as well as at high (≥ 30 mol%) dehydrogenation conversions (Based on a single passage of the reaction gas mixture through the reaction zone A). This is essentially because that with high dehydration rates in Re action zone A usually high propylene content in the product gas mixture A and, as a result, increased propylene streams in the feed gas mixture associated with the reaction zone B. The latter usually leads to increased residual oxygen flows in the (main) residual gas, which in turn are suitable for those with the increased Dehydrierumsätzen in the reaction zone A associated increased amounts of hydrogen to burn. That is, advantageously, the method of the invention includes a so self-optimizing control mechanism. With higher propane conversions are however, lower total gas loads on the catalyst feed in the reaction zone B possible. In general, it is favorable according to the invention, if the hydrogen stream in the reaction gas mixture of the reaction zone A. at the feed station for the residual gas to be recycled to the reaction zone A. in an at least stoichiometric relationship to the oxygen flow in the recycled residual gas stands.

The inventive method is applicable in a similar way when in the reaction zone B a partial ammoxidation of propene to acrylonitrile is performed. It is also applicable accordingly when in the reaction zone A the propane is replaced by iso-butane and the resulting isobutene in a corresponding manner in the reaction zone B to methacrolein and / or methacrylic acid is partially oxidized.

Comparative Examples and Examples

Comparative Example 1

A) Design of the reaction zone A (the dehydrogenation step)

The Dehydrogenation step consisted of three consecutive identical Tubular reactors, which are identical with dehydrogenation catalyst were loaded.

Of the single tube reactor was a steel tube (stainless steel of DIN material number 1.4841) of length 1300 mm, wall thickness 3.6 mm and inner diameter 53.1 mm. The Tubular reactors were from the reaction gas mixture from the top to flows through below.

At the lower end of each tubular reactor was a support rack made of the same stainless steel. The bottom of the rack was filled with the following load: 175 mm Bulk length of steatite balls (diameter 4-5 mm) of the steatite C-220 from CeramTec; 21 mm Bulk length of steatite balls (diameter 1.5-2.5 mm) of the steatite C-220 from CeramTec; 210 mm Bulk-length dehydrogenation catalyst (Pt / Sn alloy promoted with the elements Cs, K and La in oxidic form and deposited on the outer and inner surface of ZrO ₂ .SiO ₂ mixed oxide carrier strands (average length (gauss distributed in the range of 3 mm to 12 mm with maximum at about 6 mm): 6 mm, diameter: 2 mm) in elemental stoichiometry (mass ratios including support) Pt _0.3 Sn _0.6 La _3.0 Cs _0.5 K _0.2 (ZrO ₂ ) _{88 , 3} (SiO ₂ ) ₇ , (catalyst precursor preparation and activation to the active catalyst as in Example 4 of DE-A 10 219 879). 21 mm Bulk length steatite balls (diameter 1.5-2.5 mm) of the steatite C-220 from CeramTec; and finally on the remaining length of the tubular reactor again a bed of steatite (diameter 4-5 mm) of the steatite C-220 from. CeramTec.

Everyone the tube reactor was outside in the sense of a preheating line on the first 500 mm pipe length of top to bottom (towards the support back) in two, an equal distribution the supplied heat guaranteeing, Half shells made of copper (shell thickness = 200 mm) inserted, which means a heating jacket that covers it in its entirety (Horst, DE-Heidelberg, 500 mm length, 100 mm internal diameter) were electrically heated.

From down to the top, each of the tube reactors was adiabatic Track on a length of 600 mm in each case two pairs of thermally insulating acting Half shells (thickness of a half shell = 25 mm) made of MPS-Super G the Fa. Microtherm in DE, the 90 ° to each other offset one above the other attached were introduced. The insulating half shells were in turn of a cylindrical envelope of stainless steel (outer diameter = 173 mm, inner diameter = 167 mm), to which for the purpose of Heat tracing a heating jacket (length = 675 mm, inside diameter = 173 mm) Fa. Horst, DE-Heidelberg, was created. To this Way could on the adiabatic route the heat flow from the environment in the reaction tube in and out of the reaction tube in the Environment are minimized.

In each reaction tube was additionally inserted in the middle of a 1370 mm long thermal sleeve (outer diameter: 6mm, inner diameter: 4mm) into which a multiple thermocouple (from the lower reactor end upwards every 4 cm a total of 10 measuring points, thickness 3.2 mm) was introduced.

In front each individual tube reactor was one with steatite rings (made of steatite C-220 of the company CeramTec and of the geometry 7 mm × 3 mm × 3 mm = Outer diameter × diameter × height) filled steel tube the length 1300 mm switched as a heater. In it, the reaction gas mixture in each case to the inlet temperature of the subsequent tubular reactor preheated and at the same time ideally mixed. For this purpose were the heater pipes (stainless steel of DIN material number 1.4841, wall thickness 3.6 mm, inner diameter 53.1 mm) on a tube-centered length of 1200 mm by means of heating sleeves made by Horst, DE-Heidelberg electrically heated. The connection between heaters and tube reactors were thermally insulated with conventional thermal insulation materials Stainless steel pipes (stainless steel of DIN material number 1.4841, outside diameter 21.3 mm, inner diameter 16.1 mm, length approx. 700 mm).

In front the entry of the reaction gas mixture in the respective heater a supply tap was installed over the compressed air is supplied to the reaction gas mixture in each case could. Described below is the stationary operation.

the first dehydrogenation reactor, a reaction gas starting mixture was made 395 g / h crude propane (first propane-containing feed stream), 500 g / h of water and 3649 g / h (total) of residual gas as recycle gas (second propane containing feed stream) having a temperature of 400 ° C and a Pressure of 2.8 bar supplied.

The crude propane contained:

The (total) residual gas (circulating gas) contained:

The composition of the crude propane and all other gas compositions was determined by gas chromatography [HP 6890 with Chem-Station, detectors: FID, WLD, separation columns: Al ₂ O ₃ / KCL (Chrompack), Carboxene 1010 (Supelco)]. In water vapor-containing gas mixtures was this condensed in advance of the gas chromatographic analysis by cooling and optionally depressurization in a water. The uncondensed gas was analyzed and this dry calculated gas (ie, the amount of water vapor actually contained in the gas mixture to be analyzed was disregarded) is all data.

The Reaction gas starting mixture was produced in an evaporator, the was connected upstream of the first heater. The evaporator itself was also run as a heater. He was 395 g / h gaseous crude propane (65 ° C, 5 bar), 3649 g / h (total) residual gas (circulating gas) (50 ° C, 2.8 bar) and 500 g / h water (20 ° C, 3 bar) fed. The heating of the evaporator was at a gas mixture exit temperature from 200 ° C regulated. The evaporator was in the corresponding with the first heater Connected how the reactors were connected to the heaters.

The Heating of the first heater was regulated so that the evaporator in the first heater passed gas mixture the first heater with a temperature of 400 ° C left (the required wall temperature was about 440 ° C). Then, the reaction gas starting mixture led into the first tubular reactor and in the preheating section thereof to a reaction zone inlet temperature from 423.3 ° C further heated.

The Temperature of the reaction gas mixture passed through the passage the first tube reactor a maximum (called hot-spot temperature) of 549.1 ° C (The quantitative information given here refers to the Operating state after 200 operating hours; in the further course of the In operation, the different temperatures were so rebuilt that the single-pass conversion and the space-time yield remained substantially constant; in a similar way was also already in the first 200 operating hours), which in Course of continuous operation of the pilot plant due gradual Catalyst deactivation in the flow direction migrated (the migration speed was about 0.03 mm / h).

The reaction gas mixture leaving the first dehydrogenation reactor had the following contents:

His Temperature was 509.1 ° C and its pressure was 2.55 bar.

In front the entry into the subsequent heater were the reaction gas mixture 130 Nl / h of compressed air added (23 ° C, 4.2 bar).

Then the reaction gas mixture became by means of electrical heating possibilities of the heater (wall temperature about 540 ° C) and the preheating of the subsequent (second) reaction tube (wall temperature about 560 ° C) to 477 ° C (inlet 2nd reaction zone) is heated. The pressure of the reaction gas mixture at this point was 2.55 bar.

When passing through the second tubular reactor, the temperature of the reaction gas mixture passed through a maximum of 515.5 ° C, which in the course of continuous operation of the pilot plant due to gradual catalyst deactivation in the flow direction migrated (the migration rate was about 0.1 mm / h). The reaction gas mixture leaving the second dehydrogenation reactor had the following contents:

His Temperature was 493.8 ° C and its pressure was 2.5 bar.

In front the entry into the subsequent heater were the reaction gas mixture 195 Nl / h of compressed air added (23 ° C, 4.2 bar).

Then the reaction gas mixture became by means of electrical heating possibilities of the heater (wall temperature ca.480 ° C) and the preheating section of the following (third) reaction tube (wall temperature about 480 ° C) to 464 ° C (inlet 3rd reaction zone) is heated. The pressure of the reaction gas mixture at this point was 2.5 bar.

When passing through the third tubular reactor, the temperature of the reaction gas mixture went through a maximum of 492.8 ° C, which in the course of continuous operation of the pilot plant as a result of gradual catalyst deactivation in the flow direction migrated (the migration rate was about 0.1 mm / h). The reaction gas mixture leaving the third dehydrogenation reactor had the following contents:

His Temperature was 480.4 ° C and its pressure was 2.45 bar.

In order to surrendered over the dehydrogenation on the one-time passage of the starting reaction gas mixture relative total dehydrogenation of the propane of 19.83 mol%.

at advanced deactivation of dehydrogenation catalyst beds the process was interrupted and this as in DE-A 10028582 regenerated described. This was essentially always the case Case when the hotspot temperature in all three tubular reactors was about 580 ° C.

In surprising Way, the deactivation of the Dehydrierkatalysatorschüttungen at increased working pressure slower.

B) Design of the separation zone A

By direct cooling with sprayed chilled Water (T = 20 ° C) became the product gas mixture leaving the third dehydrogenation reactor (Product gas mixture A) in a direct cooler (quench) in cocurrent at 40 ° C chilled (while gaseous remaining mixture was in the opposite direction to the Produktgaseinströmrichtung led out of the water quench). About 75 wt .-% of the water vapor contained in the product gas mixture (Water vapor was the reaction gas starting mixture of the reaction zone A was added and formed in the dehydrogenation step by combustion of hydrogen as well as possibly Hydrocarbon with atmospheric oxygen; the heat of combustion stopped the reaction temperature in the reaction gas mixture of the reaction zone A largely upright) condensed in the direct cooling. The thereby forming condensate was by means of level control the water quench out and its disposal. For the rest was that for direct cooling used water is circulated (pumped), by indirect heat exchange recooled and for the purpose of direct cooling sprayed again.

Instead of the described direct cooling with water, the product gas mixture from the dehydrogenation step can also first cooled by it be that with the product gas mixture in an indirect heat exchanger (eg in a shell-and-tube heat exchanger in cocurrent or in countercurrent) the reaction gas starting mixture to be fed to the dehydrogenation stage warms up (at 400 ° C). The product gas mixture of the dehydrogenation stage cools from about 500 ° C to about 200 to 300 ° C from.

A further cooling of the product gas mixture of the dehydrogenation step can be effected in that it is used in an indirect heat exchanger to heat the starting gas mixture for the below to be described partial oxidation of propene produced in the dehydrogenation step, and / or that it is used to be described later absorber exhaust gas cooling, preferably two-stage expansion of the same by means of expansion turbines compensate or compensate by preheating.

After that the product gas mixture is at a temperature of about 180 ° C. Then you can by means of air and / or surface water cooler a temperature in the range of 30 ° C up to 60 ° C chilled become.

In the coolers collect integrated or downstream of this dropper condensed out on cooling Water and lead it was regulated for disposal.

The so cooled and relieved of water vapor, located at a pressure of about 2 bar Product gas mixture of the dehydrogenation step was subsequently to a Compressed pressure from 10 to 13 bar.

In in terms of application technology appropriate manner The compaction was carried out in two stages to high compaction temperatures to avoid (this purpose served already the previously carried out cooling, the Water vapor deposition relieves the compressor power required addition). In the first stage was compressed to a pressure of 4 to 4.5 bar. The outlet temperature of the gas mixture was on leaving the Compressor about 115 ° C.

In a downstream indirect heat exchanger (air cooler or Surface water cooler) the gas mixture is cooled again to 40 to 60 ° C, with further water vapor condensation took place. Droplets collected the condensate and kept it level out.

In the second compressor stage was based on a pressure of compressed about 4 bar to a final pressure of 10 to 13 bar. The Exit temperature when leaving the compressor was about 126 ° C.

In two further downstream indirect heat exchangers (initially a air cooler (usually a shell-and-tube heat exchanger; the to be cooled Gas flows through appropriate that Pipe interior) and then a surface water cooler) the compressed gas mixture first at 40 to 60 ° C and then to 30 ° C cooled. This condensing water was again by means of mist eliminators separated and led out. When leaving the second compressor, the gas mixture contains only 0.2% by weight of water. The low water content avoids a two-phase in the subsequent absorption and the high pressure reduces the to Purpose of absorption needed Amount of absorbent.

While large-scale turbo compressors (not oil-lubricated, dry-running, non-contact compressors) are used for the purpose of compaction (eg of the type 12 MH 4B, the company Mannesmann DEMAG, DE), diaphragm compressor MV 3459 II of the Fa. Sera used. In principle, the compressors can be driven both by electric motors and by steam or gas turbines. Often the drive via steam turbines is the most economical. The gas mixture compressed and cooled as described was fed directly to an absorption column directly above the bottom (about 4220 g / h). It had the following contents:

The Absorption column consisted of stainless steel 1.4571. The column inside diameter was 80 mm, the wall thickness was 4 mm and the column length was 1.70 m.

About 70% of the volume of the absorption column was filled with packing elements from Montz (Montz BSH-750, specific surface area 750 m ² / m ³ ) as separating internals. The packing elements followed each other directly and began at the height of 1/5 of the column length from below. The absorption column was neither cooled nor heated from the outside. At the top of the column, technical tetradecane from the company Haltermann, DE, of the type PKWF 4/7 of was used as absorbent at a charging temperature of 30 ° C. (gas chromatographic analysis by means of FID detection gave at the beginning (fresh) the following GC area% composition:
n-tridecane 7.6%,
n-tetradecane 47.3%,
n-pentadecane 7.0%,
n-hexadecane 3.2%,
n-heptadecane 14.1% and
Residual amount 20.7%;
this composition changed in the course (after about 3000 h ^-1 ) of the continuous process operation to the following values:
n-tridecane 2.6%,
n-tetradecane 39.5%,
n-pentadecane 9.4%,
n-hexadecane 4.8%,
n-Heptadecan 23.4 and
Balance 20.3%).

The irrigation density was 15 m ³ absorbent per m ^{2 of} free cross-sectional area and hours (= 25 kg / h tetradecane).

Of the from the absorption column for combustion led exhaust gas stream still contained 1700 ppm by volume of propane and 400 ppm by volume of propene. Industrially this exhaust gas flow is over an expander (eg, an expansion turbine) into the combustion guided, for the most part the compaction power used in the two-stage compaction recover and to return to the two-stage consolidation. Appropriately The expansion is also carried out in two stages to avoid unwanted condensation to avoid. The mechanical energy gained during the relaxation can either directly as a co-drive or main drive for one of the compressors and / or be used to generate electricity.

Before the relaxed absorber exhaust gas is led to the combustion, it can be industrially purpose be moderate to separate the hydrogen contained in it. This can be z. Example be done by passing the exhaust through a, usually designed to a tube, membrane which is permeable only to the molecular hydrogen. The thus separated molecular hydrogen can, for. B. in the heterogeneously catalyzed dehydrogenation or recycled (eg in fuel cells) are supplied.

in principle Hydrogen separation can also be achieved by partial condensation, Adsorption and / or rectification (preferably under pressure, eg. B. as pressure swing adsorption) can be made. Industrially it is to the usually appropriate, the absorber exhaust gas through to lead to the later to be described acid water to to narrow this down.

The absorbate contained the following contents (% by weight based on the weight of the absorbate):

Before the absorbate was passed to the top of the subsequent desorption column, it was in an indirect heat exchanger at 36 to 37 ° C heated. As a heat carrier was the liquid Drain used from the desorption column, which has a temperature of 37 ° C had.

Subsequently was the absorbate (this can be, for example, in an inverse pump or by means of a valve executed be relaxed) to a pressure of 2.5 to 3 bar (which in the case The inverse pump thereby released mechanical energy is appropriate to recompression also used in desorption column-freed absorbent) and the two-phase mixture produced at the head in the desorption column guided.

In the desorption column (inner diameter = 80 mm, length = 1.70 m, wall thickness = 4 mm) was carried out from below in countercurrent to the effluent from the Desorptionskolonnenkopf absorbant at a pressure of 3 bar air (1340 Nl / h). The amount of absorbate applied to the head was 35 l / h. The led out from the bottom of the desorption, essentially freed from desorbed components, absorbent was cooled by indirect heat exchange with absorbate, compressed to the pressure required in the absorption column, cooled by repeated indirect heat exchange (industrially expedient with surface water) to 30 ° C and then to Head of the absorption column recycled. As separation-active internals, the desorption column contained structured sheet metal packings from Montz (Montz BSH-750, specific surface area 750 m ² / m ³ ). In order to retain absorbent, the gas stream passed from the desorption column (optionally via a mechanical droplet separator) was washed with water. That is, it was passed through a packing element from Montz (Montz BSH-750, specific surface area 750 m ² / m ³ ), was added to the water (70 l / h) with a temperature of 18 ° C in countercurrent. Below the package was a catch bottom (chimney tray) attached, from which the aqueous phase could be led out. In a phase separator, it was separated into an aqueous phase and an organic phase. The very small amount of organic phase was combined with the recycled to the top of the absorption column absorbent stream. The aqueous phase was recooled and replenished with fresh water (to compensate for evaporation losses) placed back on the packing element. Washing was carried out on the desorption column.

From the wash section, the scrubbed gas stream was passed over mechanical drift eliminators (separated liquid phase is recycled to the laundry) with the following contents (if the dehydrogenation conversion is selected to be higher in the reaction zone A, the subsequent propene content may also be 8 to 12% by volume, for example the washed gas stream may also be 15 vol. % Propane, 10% by volume propene and 14.5% by volume O ₂ ):

The Temperature of the gas mixture was by indirect heat exchange at 250 ° C elevated and the gas mixture having the aforementioned composition in one Amount of 2260 Nl / l · h and an inlet pressure of 2.1 bar as a starting reaction gas mixture led into the subsequent partial oxidation.

C) Design of the reaction zone B (the partial oxidation stage)

1. First fixed bed reactor for the Step of partial oxidation of propylene (propylene) to acrolein

• Used heat exchange agent: Molten salt, consisting of 53% by weight potassium nitrate, 40 Wt .-% Nariumnitrit and 7% by weight of sodium nitrate.
• Dimension of the contact tube: 4200 mm total length, 26 mm inside diameter, 30 mm outer diameter, 2 mm wall thickness.
• Reactor: Consists of a double-walled stainless steel cylinder (cylindrical guide tube surrounded by a cylindrical outer container). The wall thicknesses were anywhere from 2 to 5 mm. The inner diameter of the outer cylinder was 168 mm. The inner diameter of the guide tube was about 60 mm. Above and below the double-walled cylinder was closed by a lid or bottom. The contact tube was guided by the cylindrical guide tube in the cylindrical container housed so that it was led out at the upper or lower end of the same (sealed) through the lid or bottom by 250 mm. The heat exchange medium was enclosed in the cylindrical container. In order to ensure as uniform as possible thermal boundary conditions on the outer wall of the contact tube over the entire contact tube length (3700 mm) in the cylindrical container, the heat exchange medium was circulated by bubbling nitrogen into the cylindrical container. By means of the rising nitrogen, the heat exchange medium was conveyed in the cylindrical guide tube from bottom to top, and then in the space between cylindrical guide tube and cylindrical outer container to flow down again (a circulation of the same quality can also by pumping (eg, propeller pumps) can be achieved). An applied to the outer jacket electrical heating, the temperature of the heat exchange medium could be controlled to the desired level. Otherwise, there was air cooling.
• Reactor charge: Viewed over the reactor, molten salt and reaction gas mixture were passed in countercurrent. The reaction gas mixture entered the top of the reactor. It was fed into the reaction tube at a temperature of 250 ° C. The salt melt entered the bottom with a temperature T ^a 320 ° C into the cylindrical guide tube and a top having a temperature T ^out of the cylindrical guide tube. The difference between T ^and T ^of was about 2 ° C. T medium = (T one + T out ) / 2.
• Contact Tube Feeding: (top to bottom) Section A: 50 cm length Pouring out Steatite rings (steatite C 220 from CeramTec) of geometry 7 mm × 7 mm × 4 mm (Outer diameter × length × inner diameter). section B: 100 cm in length Contact tube charge with a homogeneous mixture of 30 wt .-% of steatite rings (steatite C 220 from CeramTec) of the geometry 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter) and 70% by weight of full catalyst from Section C. Section C: 170 cm in length catalyst charge with annular (5 mm × 3 mm × 2 mm = outer diameter × length × inner diameter) Full catalyst according to Example 1 of DE-A 10046957. Section D: 50 cm in length Repatriation Steatite rings (steatite C 220 from CeramTec) of geometry 7 mm × 7 mm × 4 mm (Outer diameter × length × inner diameter).

2. Intercooling and Intermediate oxygen feeding

The the first fixed bed reactor leaving product gas mixture was the Purpose of intercooling (indirectly by air) through a connecting pipe (length = 400 mm, inner diameter = 26 mm, wall thickness = 2 mm, material = stainless steel) guided, that, on a length centered by 200 mm, with an inert bed out Steatite balls (steatite from CeramTec) of diameter 5 to 6 mm and directly to the Kontrahktrohr the first fixed bed reactor was flanged.

The Gas mixture entered at a temperature of more than 310 ° C in the Connecting pipe and left it with a temperature of about 140 ° C. Subsequently were the gas mixture as oxygen source 290 Nl / h to compressed Air mixed.

The resulting, mixed on a static mixer, feed gas mixture was at a temperature of 220 ° C the fixed bed reactor for the step the partial oxidation of acrolein added to acrylic acid.

3rd, second Fixed bed reactor for the step of partial oxidation of acrolein to acrylic acid

It a fixed bed reactor was used, the one for the first Step was identical. Salt melt and reaction gas mixture were over the Regarded reactor in DC. The molten salt occurred below, the reaction gas mixture also.

The contact tube charge (from bottom to top) was:
Section A: 20 cm in length
Prefill made of steatite rings (Steatite C 220 from CeramTec) of geometry 7 mm × 7 mm × 4 mm
(Outer diameter × length × inner diameter).
Section B: 100 cm in length
Catalyst charge with a homogeneous mixture of 30 wt .-% of steatite rings (steatite C 220 from. CeramTec) the geometry 7 mm × 3 mm × 4 mm
(Outer diameter × length × inner diameter) and 70 wt .-% coated catalyst from Section C.
Section C: 200 cm in length
Catalyst charging with annular (7 mm × 3 mm × 4 mm = outside diameter × length × inside diameter) coated catalyst according to Preparation Example 5 of DE-A 10046928 (at this point analogous and prepared in a corresponding manner coated catalysts can be used whose Ak however, has a stoichiometry of Mo ₁₂ V _2.8 W _1.2 Cu _2.4 O _x or of Mo ₁₂ V _3.5 W _1.3 Cu _2.4 O _x ).
Section D: 50 cm in length
Repast of steatite rings (steatite C 220 from CeramTec) of the geometry 7 mm × 7 mm × 4 mm
(Outer diameter × length × inner diameter).

The second reactor was loaded with about 3850 g / h of feed gas mixture. T ^medium is as defined for the first fixed bed reactor and was 274 ° C.

Of the Propene conversion in the first reactor was 97.7 mol% and the acrolein conversion in the second reactor was 99.4 mol%.

The contents of the product fixed-bed reactor leaving the second fixed-bed reactor at a temperature of 283 ° C. and a pressure of 1.8 bar (product gas mixture B) were:

The Catalysts used in the two reaction stages can also by the catalysts used in the examples of DE-A 10351269 be replaced. The catalyst of the first reaction stage can also by the catalysts of Examples and Comparative Examples of DE-A 10344149 be replaced. The catalyst of the second reaction stage can also by the catalysts of the examples of DE-A 10360057 and DE-A 10360058 to be replaced.

D) Design of the separation zone B (separation of the target product acrylic acid from the product gas mixture the partial oxidation)

In a direct cooler that was hot Product gas mixture (the product gas mixture B) by injecting a quench (Containing 57.4 wt .-% diphenyl ether, 20.7 wt .-% diphenyl and 20 wt .-% o-dimethyl phthalate) of the temperature 140 to 150 ° C, which also was used as an absorbent, cooled to about 180 ° C. in the lower part of the direct cooler became the quench liquid collected, taken and over a heat exchanger, in which it was recooled, sprayed back into the upper part of the quench vessel. The Product gas mixture was the direct cooler in the collected quench liquid submerged. The cooled Product gas mixture flowed then in the subsequent absorption column. Enriched in quenching high-boiling by-products, which had to be discharged. The quench circulation therefore became discontinuous once a day a subset of 1 kg of quench liquid in a container drained off discontinuously by means of a rotary evaporator was worked up. The purified absorbent became again led into the Quenchumlauf and the losses of absorbent with a mixture of diphenyl ether, Diphenyl and o-dimethyl phthalate added.

The cooled Product gas mixture was above the sump in the absorption column guided. The Absorption column was 5200 mm long, had an inner diameter of 50 mm and contained Kuehni (CH) Rombopak 9M packs. The absorption column was made of glass and had one for reasons of tempering segmented glass double jacket on. Above the first pack, viewed from below, a pressure relief valve was installed (an adjustable throttle), with the help of which a pressure loss over the Absorption column (as he is industrially would occur) from 100 to 300 mbar was simulated. The absorption column was operated with a head pressure of 1.35 bar. The glass double jacket was in five separated from each other separated segments, on the desired temperature profile imprinted has been.

From bottom to top, the segments had the following temperatures (thermostating by pumped water):
1st segment, 1020 mm long, 90-115 ° C;
2nd segment, 1250 mm long, 88 ° C;
3rd segment, 950 mm long, approx. 60 ° C;
4th segment, 950 mm long, 50-55 ° C;
5th segment, 1250 mm length, 18 ° C.

Below The topmost pack was a catcher's bottom, of which water-rich Condensate (sour water) was withdrawn. It was about one heat exchangers to 20 ° C cooled over the returned top pack in the absorption column and the excess amount (the water of reaction) discharged regulated. The exuberant Acid water was fed to a phase separator. The organic phase was over the second packing returned to the absorption column, the aqueous phase was disposed of. Immediately below the bottom of the trap was 3.5 l / h of absorption liquid with a temperature of 40 ° C given up.

The (main) residual gas discharged at the head of the absorption column (T = 18 ° C., P = 1.35 bar) had the following contents:

By indirect heat exchange the (main) residual gas became 40 ° C heated (around unwanted Exclude condensation), by means of a compressor KNF Neuberger type PM 17739-1200 (large-scale a driven by an electric motor turbo compressor, z. From the Type 12MH4B, from Mannesmann DEMAG, DE) to 2.75 bar and as a cycle gas fraction (as the first fraction of the (total) residual gas) for charging the first dehydrogenation reactor into the reaction zone A returned.

Out the bottom of the absorption column was 3.9 l / h bottoms liquid stood withdrawn regulated. A second, smaller current of about 0.5 l / h was in the direct cooler for the name is Product gas mixture returned.

Of the Sump discharge (he can first nor a stripping as described in DE-A 10336386 a partial amount of residual gas (preferably before substantially acrylic acid washed free) are subjected to; the loaded stripping gas can then as described in DE-A 10336386 in the absorption column be recycled) was placed on the head of a stripping column (50 mm internal diameter, 2000 mm in length, made of glass, double jacket for the purpose of tempering (150 ° C), 20 sieve trays (hole diameter 6.5 mm, triangular division) out. There at about 200 mbar and 150 ° C all low boilers (especially propane and propene), the main part of acrylic acid, medium boilers (eg acetic acid) as well as a portion of the absorbent. From the bottom of the stripping column Was by means of a pump (Hermetic, 500 l / h) a circulating stream (200 l / h) and over a heat exchanger (Forced circulation evaporator) in the stripping column recycled (with 190 ° C). The run at the bottom, low-load current of solvent went back to the acrylic acid absorption column and was below the bottom of the catch, where the water of reaction (Sour water) was discharged, fed again.

The stripping column escaping vapors were fed into a column free of internals, to make them rise in their condensate. The latter was taken in an amount of 100 l / h from the bottom of this column, passed through a heat exchanger in which the heat of condensation was removed, and then recycled at a temperature of 20 ° C at the top of this column.

Out the bottom of this column was further stripped 600 ml / h (contains the target product Acrylic acid, water and still residual absorbent) from which the acrylic acid at Demand can be separated in a conventional manner. At the head the column free of internals was again a sour water quench mounted, which separated by a collection tray from the column was. The condensate drawn off on the collecting tray was filled with the from the absorption column (for the product gas stream of the partial oxidation) withdrawn sour water mixed and at 18 ° C chilled returned to the sour water quench. Excess acid water it was delivered. The gas escaping from the sour-water quench was as another propane and unreacted propylene containing residual gas after compression as a further cycle gas fraction (second Proportion of (total) residual gas) to charge the first dehydrogenation reactor recycled to the reaction zone A.

Comparative Example 2

It was proceeded as in Comparative Example 1. The third dehydrogenation reactor leaving product gas mixture A was, however, through a flow divider in two halves of identical composition divided and one half in the separation zone A and the other half as part of the starting reaction gas mixture of the first dehydrogenation reactor recycled to the reaction zone A.

To the (total) residual gas was compressed to a pressure of 4 bar and with the thus compressed (total) residual gas as propellant jet one Jet pump operated, the conveying direction of a through Hover over one Mixing section and a diffuser on 3 bar relaxed propulsion jet in the first heater upstream evaporator and the Suction nozzle in the direction of the one flow half of the product mixture A pointed and the connection "suction nozzle - mixing section -Diffusor "the sole one Connection between the attributed half of the product mixture A and the first heater upstream Evaporator was.

Already after a continuous operating time of 10 h, the Dehydrogenation catalysts to lose activity.

Comparative Example 3 (all pressures are as always in this document absolute pressures, as far as nothing else explicitly mentioned becomes)

A) Design of the reaction zone A (the dehydrogenation step)

The Dehydrogenation step consisted of three consecutive identical Tubular reactors, which are identical with dehydrogenation catalyst were loaded.

Of the single tube reactor was a steel tube (stainless steel of DIN material number 1.4841) of length 1300 mm, wall thickness 3.6 mm and inner diameter 53.1 mm. The Tubular reactors were from the reaction gas mixture from the top to flows through below.

At the lower end of each tubular reactor was a support rack made of the same stainless steel. The bottom of the rack was filled with the following load: 175 mm Bulk length of steatite balls (diameter 4-5 mm) of the steatite C-220 from CeramTec; 21 mm Bulk length of steatite balls (diameter 1.5-2.5 mm) of the steatite C-220 from CeramTec; 210 mm Bulk length dehydrogenation catalyst (Pt / Sn alloy promoted with the elements Cs, K and La in oxidic form and deposited on the outer and inner surface of ZrO ₂ .SiO ₂ mixed oxide carrier strands (average length (Gauss distributed in the range of 3 mm to 12 6 mm diameter: 6 mm, diameter: 2 mm) in elemental stoichiometry (mass ratios including support) Pt _0.3 Sn _0.6 La _3.0 Cs _0.5 K _0.2 (ZrO ₂ ) _88.3 (SiO ₂ ) _7.1 (catalyst precursor preparation and activation to the active catalyst as in Example 4 of DE-A 10 219 879). 21 mm Bulk length steatite balls (diameter 1.5-2.5 mm) of the steatite C-220 from CeramTec; and finally on the remaining length of the tubular reactor again a bed of steatite (diameter 4-5 mm) of the steatite C-220 from. CeramTec.

Everyone the tube reactor was outside in the sense of a preheating line on the first 500 mm pipe length of top to bottom (towards the support back) in two, an equal distribution the supplied heat guaranteeing, Half shells made of copper (shell thickness = 200 mm) inserted, which means a heating jacket that covers it in its entirety (Horst, DE-Heidelberg, 500 mm length, 100 mm internal diameter) were electrically heated.

From down to the top, each of the tube reactors was adiabatic Track on a length of 600 mm in each case two pairs of thermally insulating acting Half shells (thickness of a half shell = 25 mm) made of MPS-Super G the Fa. Microtherm in DE, the 90 ° to each other offset one above the other attached were introduced. The insulating half shells were in turn of a cylindrical envelope of stainless steel (outer diameter = 173 mm, inner diameter = 167 mm), to which for the purpose of Heat tracing a heating jacket (length = 675 mm, inside diameter = 173 mm) Fa. Horst, DE-Heidelberg, was created.

On this way could on the adiabatic route the heat flow from the environment into the reaction tube and out of the reaction tube be minimized out into the environment.

In each reaction tube was additionally inserted in the middle of a 1370 mm long thermal sleeve (outer diameter: 6mm, inner diameter: 4mm) into which a multiple thermocouple (from the lower reactor end upwards every 4 cm a total of 10 measuring points, thickness 3.2 mm) was introduced.

In front each individual tube reactor was one with steatite rings (made of steatite C-220 of the company CeramTec and of the geometry 7 mm × 3 mm × 3 mm = Outside diameter × inside diameter × height) filled steel pipe the length 1300 mm switched as a heater. In it, the reaction gas mixture in each case to the inlet temperature of the subsequent tubular reactor preheated and at the same time ideally mixed. For this purpose were the heater pipes (stainless steel of DIN material number 1.4841, wall thickness 3.6 mm, inner diameter 53.1 mm) on a tube-centered length of 1200 mm by means of heating sleeves made by Horst, DE-Heidelberg electrically heated. The connection between heaters and tube reactors were thermally insulated with conventional thermal insulation materials Stainless steel pipes (stainless steel of DIN material number 1.4841, outside diameter 21.3 mm, inner diameter 16.1 mm, length approx. 700 mm).

In front the entry of the reaction gas mixture in the respective heater a supply tap was installed over the compressed air is supplied to the reaction gas mixture in each case could. Described below is the stationary operation.

The first dehydrogenation reactor was a reaction gas starting mixture of 300 g / h crude propane (first propane-containing feed stream), 375 g / h of water and 3768 g / h (total) residual gas as a cycle gas (two absolute propane-containing feed stream) at a temperature of 400 ° C and a pressure of 2.6 bar absolute fed (in industrial operation, the input pressure would be appropriate about 0.5 bar higher selected to the increased pressure loss (due to higher flow velocities) in reaction zone A).

The crude propane contained:

The (total) residual gas (circulating gas) contained:

The composition of the crude propane and all other gas compositions was determined by gas chromatography [HP 6890 with Chem-Station, detectors: FID, WLD, separation columns: Al ₂ O ₃ / KCL (Chrompack), Carboxene 1010 (Supelco)]. In the case of water vapor-containing gas mixtures, this gas condensate was previously condensed out by cooling and, if appropriate, decompression in a water separator. The uncondensed remaining gas was analyzed and this dry calculated gas (ie, the amount of water vapor actually contained in the gas mixture to be analyzed was disregarded) are all figures.

The Reaction gas starting mixture was produced in an evaporator, the was connected upstream of the first heater. The evaporator itself was also run as a heater. He was 300 g / h gaseous crude propane (65 ° C, 5 bar), 3768 g / h (total) residual gas (circulating gas) (50 ° C, 2.8 bar) and 375 g / h of water (20 ° C, 3 bar) fed. The heating of the evaporator was at a gas mixture exit temperature from 200 ° C regulated. The evaporator was in the corresponding with the first heater Connected how the reactors were connected to the heaters.

The Heating of the first heater was regulated so that the evaporator in the first heater passed gas mixture the first heater with a temperature of 400 ° C left (the required wall temperature was about 440 ° C). Then, the reaction gas starting mixture led into the first tubular reactor and in the preheating section thereof to a reaction zone inlet temperature of 460 ° C further heated.

The Temperature of the reaction gas mixture passed through the passage the first tube reactor a maximum (hot-spot temperature called) of 549.1 ° C (The quantitative information given here refers to the Operating state after 200 operating hours; in the further course of the In operation, the different temperatures were so rebuilt that the single-pass conversion and the space-time yield remained substantially constant; in a similar way was also already in the first 200 operating hours), which in Course of continuous operation of the pilot plant due gradual Catalyst deactivation in the flow direction migrated (the migration speed was about 0.03 mm / h).

The reaction gas mixture leaving the first dehydrogenation reactor had the following contents:

His Temperature was 509 ° C and his pressure was about 2.56 bar.

In front the entry into the subsequent heater were the reaction gas mixture 80 Nl / h of compressed air added (23 ° C, 4.2 bar).

Then the reaction gas mixture became by means of electrical heating possibilities of the heater (wall temperature about 540 ° C) and the preheating of the subsequent (second) reaction tube (wall temperature about 560 ° C) to 465 ° C (inlet 2nd reaction zone) is heated. The pressure of the reaction gas mixture at this point was 2.56 bar.

During the passage through the second tubular reactor, the temperature of the reaction gas mixture passed through a maximum of about 560 ° C, which in the course of the continuous operation of the pilot plant due to gradual catalyst deactivation in the flow direction migrated (the migration rate was about 0.1mm / h). The reaction gas mixture leaving the second dehydrogenation reactor had the following contents:

His Temperature was about 493 ° C and his pressure was about 2.52 bar.

In front the entry into the subsequent heater were the reaction gas mixture 98 Nl / h of compressed air added (23 ° C, 4.2 bar).

Then the reaction gas mixture became by means of electrical heating possibilities of the heater (wall temperature about 540 ° C) and the preheating of the subsequent (third) reaction tube (wall temperature about 540 ° C) to 521 ° C (inlet 3rd reaction zone) is heated. The pressure of the reaction gas mixture at this point was 2.52 bar.

When passing through the third tubular reactor, the temperature of the reaction gas mixture passed through a maximum of 570 ° C, which in the course of the continuous operation of the pilot plant as a result of gradual catalyst deactivation in the flow direction migrated (the migration rate was about 0.1 mm / h). The reaction gas mixture leaving the third dehydrogenation reactor had the following contents:

His Temperature was 480.4 ° C and its pressure was 2.48 bar.

In order to surrendered over the dehydrogenation on the one-time passage of the starting reaction gas mixture relative total dehydrogenation of the propane of 19.91 mol%.

at advanced deactivation of dehydrogenation catalyst beds the process was interrupted and this as in DE-A 10028582 regenerated described. This was essentially always the case Case when the hotspot temperature in all three tubular reactors was about 580 ° C.

In surprising Way, the deactivation of the Dehydrierkatalysatorschüttungen at increased working pressure slower.

B) Design of the separation zone A

By direct cooling with sprayed chilled Water (T = 20 ° C) became the product gas mixture leaving the third dehydrogenation reactor (Product gas mixture A) in a direct cooler (quench) in cocurrent at 40 ° C chilled (while gaseous remaining mixture was in the opposite direction to the Produktgaseinströmrichtung led out of the water quench). About 75 wt .-% of the water vapor contained in the product gas mixture (Water vapor was the reaction gas starting mixture of the reaction zone A was added and formed in the dehydrogenation step by combustion of hydrogen as well as possibly Hydrocarbon with atmospheric oxygen; the heat of combustion stopped the reaction temperature in the reaction gas mixture of the reaction zone A largely upright) condensed in the direct cooling. The thereby forming condensate was by means of level control the water quench out and its disposal. For the rest was that for direct cooling used water is circulated (pumped), by indirect heat exchange recooled and for the purpose of direct cooling sprayed again.

Instead of the described direct cooling with water, the product gas mixture from the dehydrogenation stage can also first be cooled by warming up the reaction gas starting mixture to be fed to the dehydrogenation stage with the product gas mixture in an indirect heat exchanger (eg in a tube bundle heat exchanger in cocurrent or countercurrent heat) B. to a temperature in the range of 350 to 450 ° C). The product gas mixture of the dehydrogenation stage cools to a corresponding extent of the high outlet temperature of the dehydrogenation stage (eg from 450 to 550 ° C) to about 200 to 300 ° C from.

A further cooling the product gas mixture of the dehydrogenation stage can be effected by that it is in an indirect heat exchanger is used, the starting gas mixture for the below to be described To heat partial oxidation of the propene produced in the dehydrogenation step, and / or that it is used in the later to be described absorber exhaust gas cooling, preferably two-stage, expanding the same by means of expansion turbines compensate or compensate by preheating.

After that the product gas mixture is at a temperature of about 180 ° C. Then you can by means of air and / or surface water cooler a temperature in the range of 30 ° C up to 60 ° C chilled become.

In the coolers collect integrated or downstream of this dropper condensed out on cooling Water and lead it was regulated for disposal.

The so cooled and relieved of water vapor, located at a pressure of about 2 bar Product gas mixture of the dehydrogenation step was subsequently to a Compressed pressure from 10 to 13 bar.

In in terms of application technology appropriate manner The compaction was carried out in two stages to high compaction temperatures to avoid (this purpose served already the previously carried out cooling, the Water vapor deposition relieves the compressor power required addition). In the first stage was compressed to a pressure of 4 to 4.5 bar. The outlet temperature of the gas mixture was on leaving the Compressor about 115 ° C.

In a downstream indirect heat exchanger (air cooler or Surface water cooler) the gas mixture is cooled again to 40 to 60 ° C, with further water vapor condensation took place. Droplet collect the condensate and pass it was leveled.

In the second compressor stage was based on a pressure of compressed about 4 bar to a final pressure of 10 bar (here may optionally be compressed to a pressure of up to 13 bar and more). The exit temperature when leaving the compressor was about 126 ° C.

In two further downstream indirect heat exchangers (initially a air cooler (usually a shell-and-tube heat exchanger; the to be cooled Gas flows through appropriate that Pipe interior) and then a surface water cooler) the compressed gas mixture first at 40 to 60 ° C and then to 30 ° C cooled. This condensing water was again by means of mist eliminators separated and brought out. When leaving the second compressor, the gas mixture contains only about 0.2 wt .-% water. The low water content reduces the water attack and thus avoids by a two-phase of the liquid conditional operational problems in the subsequent absorption and the high pressure reduces the amount needed for the purpose of absorption on absorbent.

While large-scale turbo compressors (not oil-lubricated, dry-running, non-contact compressors) are used for the purpose of compaction (eg of the type 12 MH 4B, the company Mannesmann DEMAG, DE), diaphragm compressor MV 3459 II of the Fa. Sera used. In principle, the compressors can be driven both by electric motors and by steam or gas turbines. Often the drive via steam turbines is the most economical. The compressed and cooled as described gas mixture was fed directly above the bottom of an absorption column (about 4650 g / h). It had the following contents:

The Absorption column consisted of stainless steel 1.4571. The column inside diameter was 80 mm, the wall thickness was 4 mm and the column length was 1.70 m.

About 70% of the volume of the absorption column was filled with packing elements from Montz (Montz BSH-750, specific surface area 750 m ² / m ³ ) as separating internals. The packing elements followed each other directly and began at the height of 1/5 of the column length from below. The absorption column was neither cooled nor heated from the outside. At the top of the column technical tetradecane from the company Haltermann, DE, of the type PKWF 4/7 af was used as absorption medium at a charging temperature of 30 ° C. (gas chromatographic analysis by means of FID detection gave at the beginning (fresh) the following GC area% composition:
n-tridecane 7.6%,
n-tetradecane 47.3%,
n-pentadecane 7.0%,
n-hexadecane 3.2%,
n-heptadecane 14.1% and
Residual amount 20.7%;
this composition changed in the course (after about 3000 h ^-1 ) of the continuous process operation to the following values:
n-tridecane 2.6%,
n-tetradecane 39.5%,
n-pentadecane 9.4%,
n-hexadecane 4.8%,
n-Heptadecan 23.4 and
Balance 20.3%).

The irrigation density was 15 m ³ absorbent per m ^{2 of} free cross-sectional area and hours (= 28 kg / h tetradecane).

Of the from the absorption column for combustion led exhaust gas stream still contained 950 ppm by volume of propane and 250 ppm by volume of propene. On a large scale, this exhaust gas flow over a Expander (eg an expansion turbine) led into the combustion to the majority the compaction power used in the two-stage compaction recover and to return to the two-stage consolidation. Appropriately The expansion is also carried out in two stages to avoid unwanted condensation to avoid. The mechanical energy gained during the relaxation can either directly as a co-drive or main drive for one of the compressors and / or be used to generate electricity.

Before the relaxed absorber exhaust gas is led to the combustion, it can be industrially purpose be moderate to separate the hydrogen contained in it. This can be z. Example be done by passing the exhaust through a, usually designed to a tube, membrane which is permeable only to the molecular hydrogen. The thus separated molecular hydrogen can, for. B. in the heterogeneously catalyzed dehydrogenation or recycled (eg in fuel cells) are supplied.

in principle Hydrogen separation can also be achieved by partial condensation, Adsorption and / or rectification (preferably under pressure) made become. Industrially it is to the usually appropriate, the absorber exhaust gas through to lead to the later to be described acid water to this constrict.

The absorbate contained the following contents (% by weight based on the weight of the absorbate):

Before the absorbate was passed to the top of the subsequent desorption column, it was in an indirect heat exchanger to 60 ° C heated.

Subsequently was the absorbate (this can be, for example, in an inverse pump or by means of a valve executed be relaxed) to a pressure of 2.7 bar (which in the case of the inverse Pump thereby released mechanical energy is expedient for recompression also used in desorption column-freed absorbent) and the two-phase mixture produced at the head passed into the desorption column.

Into the desorption column (inner diameter = 80 mm, length = 1.70 m, wall thickness = 4 mm) was carried out from below in countercurrent to the effluent from the Desorptionskolonnenkopf absorbant at a pressure of 3 bar air (1269 Nl / h). The amount of absorbate applied to the head was 33.7 l / h. The brought out from the bottom of the desorption, essentially freed from desorbed components, absorbent was cooled by indirect heat exchange, compressed to the pressure required in the absorption column, cooled by repeated indirect heat exchange (industrially expedient with surface water) to 16 ° C and then to the head of Recirculated absorption column. As separation-active internals, the desorption column contained structured sheet metal packings from Montz (Montz BSH-750, specific surface area 750 m ² / m ³ ). In order to retain absorbent, the gas stream passed from the desorption column (optionally via a mechanical droplet separator) was washed with water. That is to say, it was passed through a packing element from Montz (Montz BSH-750, specific surface area 750 m ² / m ³ ), to which water (70 l / h) with a temperature of 18 ° C. was charged in countercurrent. Below the package was a catch bottom (chimney tray) attached, from which the aqueous phase could be led out. In a phase separator, it was separated into an aqueous phase and an organic phase. The very small amount of organic phase was combined with the recycled to the top of the absorption column absorbent stream. The aqueous phase was recooled and replenished with fresh water (to compensate for evaporation losses) placed back on the packing element. Washing was carried out on the desorption column.

From the washing section, the washed gas stream was taken out via mechanical droplet separators (separated liquid phase is returned to the laundry) with the following contents (If the dehydrogenation conversion is selected to be higher in the reaction zone A, the subsequent propene content may also be 8 to 12% by volume, for example the washed gas stream may also contain 15% by volume of propane, 10% by volume of propene and 14.5% by volume % O ₂ ):

The Temperature of the gas mixture was by indirect heat exchange at 250 ° C elevated and the gas mixture having the aforementioned composition in one Amount of 2128 Nl / l · h and an inlet pressure of 2.1 bar as a starting reaction gas mixture led into the subsequent partial oxidation.

C) Design of the reaction zone B (the partial oxidation stage)

1. First fixed bed reactor for the Step of partial oxidation of propylene (propylene) to acrolein

• Used heat exchange agent: Molten salt, consisting of 53% by weight potassium nitrate, 40 Wt .-% Nariumnitrit and 7% by weight of sodium nitrate.
• Dimension of the contact tube: 4200 mm total length, 26 mm inside diameter, 30 mm outer diameter, 2 mm wall thickness.
• Reactor: Consists of a double-walled stainless steel cylinder (cylindrical guide tube surrounded by a cylindrical outer container). The wall thicknesses were anywhere from 2 to 5 mm. The inner diameter of the outer cylinder was 168 mm. The inner diameter of the guide tube was about 60 mm. Above and below the double-walled cylinder was closed by a lid or bottom. The contact tube was guided by the guide tube in the cylindrical container housed so that it was led out at the upper or lower end of the same (sealed) through the lid or bottom by 250 mm. The heat exchange medium was enclosed in the cylindrical container. In order to ensure as uniform as possible thermal boundary conditions on the outer wall of the contact tube over the entire contact tube length (3700 mm) in the cylindrical container, the heat exchange medium was circulated by bubbling nitrogen into the cylindrical container. By means of the rising nitrogen, the heat exchange medium was conveyed in the cylindrical guide tube from bottom to top, and then in the space between the cylindrical guide tube and cylindrical outer container back down to flow (a circulation of the same quality can be achieved by pumping (eg., Propeller pumps) who the ). By an applied to the outer jacket electric heater, the temperature of the heat exchange medium could be controlled to the desired level. Furthermore consisted of air cooling. Reactor charge: Viewed over the reactor, molten salt and reaction gas mixture were passed in countercurrent. The reaction gas mixture entered the top of the reactor. It was fed into the reaction tube at a temperature of 250 ° C.
• The salt melt entered the bottom with a temperature T ^a 335 ° C into the cylindrical guide tube and a top having a temperature T ^out of the cylindrical guide tube. The difference between T ^and T ^of was about 2 ° C. T medium = (T one + T out ) / 2.
• Contact Tube Feeding: (top to bottom) Section A: 50 cm length Pouring out Steatite rings (steatite C 220 from CeramTec) of geometry 7 mm × 7 mm × 4 mm (Outer diameter × length × inner diameter). section B: 100 cm in length Contact tube charge with a homogeneous mixture of 30 wt .-% of steatite rings (steatite C 220 from CeramTec) of the geometry 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter) and 70% by weight of full catalyst from Section C. Section C: 170 cm in length catalyst charge with annular (5 mm × 3 mm × 2 mm = outer diameter × length × inner diameter) Full catalyst according to Example 1 of DE-A 10046957. Section D: 50 cm in length Repatriation Steatite rings (steatite C 220 from CeramTec) of geometry 7 mm × 7 mm × 4 mm (Outer diameter × length × inner diameter).

2. Intercooling and Intermediate oxygen feeding

The the first fixed bed reactor leaving product gas mixture was the Purpose of intercooling (indirectly by air) through a connecting pipe (length = 400 mm, inner diameter = 26 mm, wall thickness = 2 mm, material = stainless steel) guided, that, on a length centered by 200 mm, with an inert bed out Steatite balls (steatite of the Fa CeramTec) of diameter 5 to 6 mm and directly to the Kontrahktrohr the first fixed bed reactor was flanged.

The Gas mixture entered at a temperature of more than 310 ° C in the Connecting pipe and left it with a temperature of about 140 ° C. Subsequently were the gas mixture as an oxygen source 269 Nl / h to compressed Air mixed.

The resulting, mixed on a static mixer, feed gas mixture was at a temperature of 220 ° C the fixed bed reactor for the step the partial oxidation of acrolein added to acrylic acid.

3rd, second Fixed bed reactor for the step of partial oxidation of acrolein to acrylic acid

It a fixed bed reactor was used, the one for the first Step was identical. Salt melt and reaction gas mixture were over the Regarded reactor in DC. The molten salt occurred below, the reaction gas mixture also.

The contact tube charge (from bottom to top) was:
Section A: 20 cm in length
Prefill made of steatite rings (Steatite C 220 from CeramTec) of the geometry 7 mm × 7 mm × 4 mm
(Outer diameter × length × inner diameter).
Section B: 100 cm in length
Catalyst charge with a homogeneous mixture of 30 wt .-% of steatite rings (steatite C 220 from. CeramTec) the geometry 7mm × 3mm × 4mm
(Outer diameter × length × inner diameter) and 70 wt .-% coated catalyst from Section C.
Section C: 200 cm in length
Catalyst charge with annular (7 mm × 3 mm × 4 mm = outside diameter × length × inside diameter) coated catalyst according to Preparation Example 5 of DE-A 10046928 (at this point also analog and prepared in a corresponding manner coated catalysts can be used, the active material, however, a stoichiometry of Mo ₁₂ V _2.8 W _1.2 Cu _2.4 O _x or of Mo ₁₂ V _3.5 W _1.3 Cu _2.4 O _x ).
Section D: 50 cm in length
Repast of steatite rings (steatite C 220 from CeramTec) of the geometry 7 mm × 7 mm × 4 mm
(Outer diameter × length × inner diameter).

The second reactor was charged with approximately 3607 g / h of feed gas mixture. T ^medium is as defined for the first fixed bed reactor and was 284 ° C.

Of the Propene conversion in the first reactor was 98.0 mol% and the acrolein conversion in the second reactor was about 99.8 mol%.

The contents of the product gas mixture leaving the second fixed bed reactor at a temperature of 281 ° C. and a pressure of 1.8 bar (product gas mixture B) were:

The catalysts used in the two reaction stages can also be replaced by the catalysts used in the examples of DE-A 10351269. The catalyst of the first reaction stage can also be replaced by the catalysts of the examples and comparative examples of DE-A 10344149. The catalyst of the second reaction stage can also be replaced by the catalysts of the examples of DE-A 10360057 and DE-A 10360058. Furthermore, the partial oxidation can be carried out as a high-load process, as described in DE-A 10313210, DE-A 10313213, DE-A 10313212, DE-A 10313211, DE-A 10313208, DE-A 10313209 and in US Pat DE-A 10313214 and the appreciated in these publications State of the art is described.

D) Design of the separation zone B (separation of the target product acrylic acid from the product gas mixture the partial oxidation)

In a direct cooler, the hot product gas mixture stream (the product gas mixture B) was cocurrently charged (after combining with the loaded stripping gas stream described below) by injecting a quench liquid (containing 57.4% by weight of diphenyl ether, 20.7% by weight of diphenyl, 20 Wt .-% o-dimethyl phthalate (DMP) and as a residual amount up to 100 wt .-% phenothiazine (up to 1 wt .-%) and high boilers (eg., Acrylsäureoligomere)) of the temperature 140 to 150 ° C, which also was used as an absorbent, cooled to about 180 ° C. In the lower part of the direct cooler, the unevaporated quench liquid was collected, removed and sprayed back into the upper part of the quench vessel via a heat exchanger in which it was brought back to flow temperature. The product gas mixture stream B (combined with the loaded stripping gas) and the quench liquid atomized to jets were fed in parallel to the direct cooler. The cooled product gas mixture stream B then flowed into the subsequent absorption column. In the quench circulation, high-boiling by-products accumulated, which had to be discharged. Therefore, once a day, a portion of 1 kg of quench liquid was discharged from the quench circulation into a container which was worked up discontinuously with the aid of a rotary evaporator (8 mbar, 180 ° C.). The purified absorbent condensate obtained in this way was fed to the buffer vessel described below for the bottoms effluent from the residual gas stainless steel pressure wash column and once a week the losses of absorbent of about 2.9 g / h incurred by working up in a rotary evaporator by addition of fresh Mixture consisting of diphenyl ether (21.2% by weight), diphenyl (58.8 wt .-%) (both together = Diphyl ^® ) and o-dimethyl phthalate (20 wt .-%) replaced the condensate.

The cooled Product gas mixture B was above the bottom and below the bottom led first pack in the absorption column. The absorption column (without sump) was 5420 mm long, had an inside diameter of 50 mm and contained a total of 3.3 m Kuehni (CH) Rhombopak 9M length Packs.

The Absorption column was made of glass and dismissed for the sake of Tempering a segmented glass double jacket on. immediate above the second pack, viewed from below, was a pressure decoupling valve installed (an adjustable throttle), with the help of which a pressure loss over the Absorption column (as he is industrially would occur) from 100 to 300 mbar could be simulated. The absorption column was with a head pressure of 1.30 bar and a sump pressure of Operated 1.40 bar. The bottom temperature was about 113 ° C. The glass double jacket was in five subdivided successive segments operated separately over the the absorption column was impressed on the subsequent temperature profile. This segmentation was also followed by the internal structure of the absorption column.

The segments had the following temperatures from bottom to top (their thermostatting was carried out in each case by means of circulated water, or in the 1st segment with heat transfer oil of the corresponding temperature) and were designed as follows:
1st segment: 1020 mm in length (beginning at the feed point of the product gas mixture B beginning), 113 ° C, 0.4 m rhombopak between the immediately above the bottom of the absorption column feed point of the product gas mixture B and the feed point of from the bottom of the column continuously withdrawn and rezierkulierter in the column (at about 113 ° C, in an amount of about 200 l / h) bottoms liquid and 0.6 m rhombopak above this Aufgabestelle.
The distance between the two rhombopaks was about 2 cm.
2nd segment: 1250 mm length, 75 ° C., at the lower end was the pressure decoupling valve and above the pressure decoupling valve to the feed point of the low boiler condensate stream from the purifying distillation described below was 0.5 m rhombopak.
3rd segment: 950 mm long, 60 ° C, 0.6 m Rhombopak, starting just above the low boiling condensate feed point of the 2nd segment.
4th segment: 950 mm long, 50-55 ° C, 0.6 m rhombopak, which terminate immediately below the feed point of the main absorbent stream described below.
5th Segment: 1250 mm in length, 20 ° C, 0.6m Rhombopak between the immediately above the Absorptionsmittelaufgabe in the 4th segment for the expiration of the sour water (described below) attached catch bottom and the point of charge of recirculated sour water at the top of the absorption column.

From the above in the 5th segment catch ground was continuously rich in water Kon Densate (sour water) deducted (about 70.2 l / h). It was cooled indirectly via a heat exchanger to 20 ° C and returned above the top rhombopak packing back into the absorption column (70 liters / h). The non-recirculated portion of the withdrawn sour water (the water of reaction) was discharged under controlled conditions and fed to an acid water extraction stage. In this, the discharged sour water was combined in an ambient temperature glass stirred vessel (1 liter volume, 80 mm diameter) at ambient pressure with the smaller partial stream of (described below) lightly loaded (with acrylic acid and minor components) absorbent and mixed. The resulting mixture was continuously drained into a glass phase separation vessel (7.5 liters in volume, 150 mm in diameter) also maintained at ambient temperature by free overflow. There, the overflowed liquid mixture separated into two liquid phases, with short-chain acidic secondary components (eg, acetic acid) preferably in the specific lighter aqueous phase and acrylic acid preferably in the specific heavier organic phase. The organic phase is combined with the larger partial flow of the low-loaded absorbent by means of a positive-promoting Diaphragm metering and fed this total flow (about 3.0 l / h, 40 ° C) as already described directly below the acid water trap bottom as the main absorbent stream of the absorption column.

The (main) residual gas leaving the absorption column at the top (T = 20 ° C., P = 1.30 bar) had the following contents:

By indirect heat exchange the (main) residual gas became 40 ° C heated (around unwanted Exclude condensation), by means of a membrane compressor (on an industrial scale by means of an electric motor driven turbocompressor, z. B. the type 12MH4B, the company Mannesmann DEMAG, DE) to 2.70 bar and about 78% by volume as cycle gas (Total) residual gas to feed the first dehydrogenation reactor in the reaction zone A recycled.

About 22 vol .-% of the compressed (main) residual gas were abge with relaxation to 1.6 branches and in a stainless steel pressure wash column (material 1.4571, inner diameter 30 mm, filling: 2 m Rhombopak 9M) at about 1.6 bar and about 51 ° C with the total described below total flow of least loaded absorbent in countercurrent (The least loaded absorbent was added at the top and the gas passed down the column).

In another stainless steel pressure column (material 1.4571, inner diameter 30 mm, filling: 3 m Rhombopak 9M, double jacket tempered with oil = stripping column) This washed gas stream was at about 1.5 bar and about 119 to Used 125 ° C, from the loaded absorbent stream, which from the sump the absorption column was withdrawn continuously (about 3.5 l / h, about 113 ° C), to drain out low boiling components (eg as in DE-A 10336386). The resulting with low-boiling components (more easily as acrylic acid) loaded gas stream (loaded stripping gas) was tempered at 170 ° C. Double sheathed cable (diameter 15 mm, outer sheath stainless steel corrugated hose) heated and at the entrance of the quench vessel with the hot product gas mixture stream B united.

At the bottom of the stainless steel pressure wash column became the absorbent stream slightly loaded was regulated in the already mentioned buffer container (glass, 5 liters of volume) (this container is, as already described, also passed the obtained in a rotary evaporator absorbent condensate).

from buffer tank was by means of a membrane pump a larger partial flow of 2.5 l / h passed to the absorption column at low loaded absorbent and there as already described as part of the main absorbent stream immediately below the catch bottom for the discharge of acid water in the Segment 4 abandoned. The smaller partial flow of 460 ml / h at low loaded absorbent was, as also described, by means of another membrane pump into the glass stirring vessel of the acid water extraction stage directed.

Out the bottom circulation (about 100 l / h, compressed air membrane pump) of the stripping column were about 3.5 kg bottoms liquid over a Stainless steel fabric filter and a control valve was withdrawn regulated and for the purpose of purifying the vacuum distillation unit fed.

The Vacuum distillation unit consisted of a mirrored and vacuum insulated Glass column (pure column) with 50 mm inner diameter and 4000 mm Length. With forced exhaust sump circulation (approx. 250 l / h, peripheral centrifugal pump) a sump temperature of 191 ° C was maintained (p = 4 bar). Of the Absolute pressure in the sump was about 230 mbar, the head pressure was at 100 mbar. For the purpose of polymerization inhibition were above the sump mirror 52 Nl / h air supplied.

Between the bottom of the vacuum distillation column and the inflow of the product laden Stream from the bottom circulation of the stripping column in the vacuum distillation column were first 6 bubble trays (ground clearance: 5 cm) and above another 15 bubble trays (Ground clearance: 5 cm) attached, above which the possibility Sampling by means of a miniature diaphragm metering existed.

Above this sampling point was 10 sieve trays (6 equidistant holes with a diameter of 6.5 mm) (ground clearance: 5 cm) attached extended up to a catching ground, of which continuous about 364 g / h of purified acrylic acid discharged as Zielpodukt and stored after cooling in a storage container were.

Secondary component contents of the target product were:

On the top of the screen trays located below the Zielproduktabzug was a further amount (732 ml / h) of discharged from the bottom of the capture acrylic acid Retaining its withdrawal temperature of about 75 ° C in the distillation column recycled.

Above the target product output, there were another 10 sieve trays (6 holes each with a diameter of 5.5 mm per tray, ground clearance: 5 cm), which allowed to enrich any remaining low boilers to the top of the column. Above this sieve trays, a further trapping tray was mounted in order to collect the low-boiling condensate, which is produced by indirect cooling in the top of the column (26 ° C., 100 mbar absolute pressure). The temperature on the bottom of the trap was 73 ° C. The low-boiling condensate contained:

From was taken from the bottom of the collecting low-boiling condensate a main flow of 570 ml / h below the low boiling condensate tray as a return fed back into the distillation column. The remaining one Low boiling condensate residual stream of 190 ml / h was cooled to 40 ° C and above the 2nd segment of the acrylic acid absorption column fed. For inhibitory wall wetting was in the uppermost region of the pure column head one with 0.5% by weight Phenothiazine stabilized target product flow of 51 ml / h with the Temperature 25 ° C over a 4-hole full jet injected.

The withdrawn at the head of the pure column by means of a membrane vacuum gas stream consisted mainly of inert gases and low boilers. In a cold trap cooled to 8 ° C., it was possible to separate 4.6 g / h of condensable low-boiling residual components from it. This liquid separated residual component condensate contained:

The "inert gas" stream remaining after the residual component condensate had the following composition:

He was above the pressure decompression valve and below it Rhomopak packing recycled to the 2nd segment of the absorption column.

At the Bottom of the pure column was regulated by a large extent that of acrylic acid liberated absorbents withdrawn and as the lowest loaded Absorptionsmittelgesamtstrom directed to the already described stainless steel pressure wash column, the pressure increase the sump circulation pump of the pure column for discharge from the vacuum was used.

The least loaded absorbent contained:

to Reduction of Eduktverlusten and to achieve a high yield were all loaded with reactant or target product analysis gas flows in one glass container (0.5 liters) merged and by means of a small diaphragm pump of the first compressor stage of Separation zone A supplied and before the local compressor with the cooled product gas mixture of Dehydrogenation combined and then compressed. Related to used propane was thus a yield of acrylic acid of 78.9 mol% achieved.

The as the target product obtained crude acrylic acid can as in the Scriptures EP-A 616998 (in accordance with the teaching of EP-A 912486) or as in the document DE-A 19606877 (the mother liquor recycling can be incorporated into the absorption and / or into the clean column) or as in the Scriptures EP-A 648732 described further purified by crystallization or rectification to pure acrylic acid which are then superabsorbent for producing water Polymerized in a conventional manner free-radically could be. Both the recovered crude acrylic acid and the won Glacial acrylic acid are outstandingly suitable for the production of esters of Acrylic acid, z. B. for the production of alkyl acrylates.

Comparative Example 4

Similar to Comparative Examples 1 and 3, in a reaction zone A together with a downstream separation zone A, a reaction gas starting mixture for a reaction zone B is produced, which has the following contents:

The reaction zone B consists of two parallel operated tandem reactor systems (reactor lines), each of which, like the tandem reactor system, is constructed from the exemplary embodiment of DE-A 103 51 269 and charged with catalyst. Each of the two reaction lines is charged with 61022 Nm ³ / h of the aforementioned reaction gas starting mixture. The intermediate air feed between the two partial oxidation stages per road 11200 Nm ³ / h. In both partial oxidation stages, reaction gas and molten salt are considered to flow countercurrently over the reactor.

The Load of the first reaction stage with propylene is about 110 Nl / lh. The inlet temperature of the salt bath in the first reaction stage is 346 ° C, the Exit temperature of the salt bath from the first reaction stage is 349 ° C. The inlet temperature of the salt bath in the second reaction stage is 272 ° C. The outlet temperature of the Salt bath from the second reaction stage is 275 ° C. The reaction gas starting mixture is fed to the first stage reactor at 300 ° C. The gas inlet temperature in the second reaction stage is 230 ° C. The salt melt pump performances are as indicated in DE-A 103 51 269. The inlet pressure in the first reaction stage is 1.9 bar. The inlet pressure into the second reaction stage is 1.4 bar. The propylene conversion in the first reaction stage is 97 mol% (related to a single pass). The acrolein turnover in the second Reaction level is 99.3 mol% (based on single pass).

The two product gas mixtures leave the two second partial oxidation stages at a temperature of 275 ° C. They are combined to a total flow of 145000 Nm ³ / h, the following, based on its total, contents, has: nitrogen 42.75% by weight, oxygen 3.0% by weight, carbon oxides 3.11% by weight, water 4.87% by weight, acrylic acid 10.84% by weight, acetic acid 0.30 wt.%, acrolein 0.06 wt.%, Diphyl 0% by weight, formic acid 0.03 wt.%, formaldehyde 0.09 wt.%, propionic 0.004 wt.%, furfural 0.003 wt.%, allyl 0.001% by weight benzaldehyde 0.009 wt.%, maleic anhydride 0.127% by weight, benzoic acid 0.015% by weight, methacrylic acid 0.01% by weight, methacrolein 0.015% by weight, phthalic anhydride 0.001% by weight, methane 0.001% by weight, hydrogen 0% by weight, Ethan 0.17% by weight, ethene 0.03% by weight, Butane 0.47% by weight, butenes 0.14% by weight, propane 33.8% by weight and propene 0.14 wt .-%.

Cooling of the product gas mixture B

The Product gas mixture B is in a direct cooler of conventional design with a diameter of 2.6 m and a height led by 14.5 m, made of austenitic steel (material 1.4571). By partial evaporation of the absorbent, the sump of the Absorption column is removed, the product gas mixture a temperature of 165.5 ° C cooled. The cooling medium is doing in cocurrent to the cooled product gas mixture B guided and over a baffle plate nozzle abandoned in the direct capacitor.

The cooled Product gas mixture B is used together with the non-vaporized absorbent in led the bottom of the subsequent absorption unit.

absorbing unit

The Absorption unit is a tray column with 45 valve trays (where the valve covers of the individual valves are not mounted) and 3 fireplace floors accomplished and is operated with a head pressure of 1.2 bar. A version with sieve trays or rain screen bottoms as separation-effective internals in the tray column is also conceivable.

Of the lowest floor of the absorption column is as a double-walled fireplace floor executed. Below this bottom, that leaving the product gas direct cooling becomes Gas mixture from in the cooler not evaporated absorbent separated.

The the direct cooling leaving mixture of non-vaporized absorbent and Product gas mixture is with a tangential pulse in the absorption column given up. A cyclone ring prevents a direct drift in the chimneys of the bottom floor.

The sump temperature of the absorption unit is 162.7 ° C. From the bottom of the absorption unit 1067 m ³ / h bottom liquid are withdrawn with the following composition: Diphyl (mixture of diphenyl and diphenyl ether in a weight ratio of 2.8: 1) 54.27% by weight, dimethyl 36.97% by weight, diacrylate 0.95% by weight, acrylic acid 4.75% by weight, phenothiazine 0.65% by weight, phthalic anhydride 0.1% by weight, benzoic acid 1.87% by weight, maleic anhydride 0.24% by weight, benzaldehyde 0.11 wt.%, formic acid 0.005 wt.%, acrolein 0.002% by weight, water 0.01% by weight, acetic acid 0.04 wt .-% and furfural 0.003 wt .-%.

The bottoms liquid (Liquid drain the absorption unit) becomes predominant as a coolant with the help of a centrifugal pump in the direct cooling of the product gas mixture B returned. A Subset of 0.3% based on the total amount is taken and fed to a distillation unit.

Of the Inlet of fresh or slightly loaded absorbent in the absorption column is counted on the 34th floor from below, immediately below the 3rd fireplace ground. This intake of lightly loaded Absorbent consists of a polymerization inhibitor-containing Partial stream of 86.6 wt .-% based on the total feed, the from that to be described below used for cycle gas scrubbing Wash unit emerges, and from a partial flow of 13.4 wt .-% based on the total feed from the sour water extraction. The temperature of the feed stream is 51.9 ° C.

From the lowest bottom of the absorption unit, which is designed as a catch bottom (chimney tray), 1728 m ³ / h of the absorbent laden with acrylic acid (the absorbate) are taken (as a further liquid effluent of the absorption column). The removal temperature is 119.6 ° C. The absorbed via a centrifugal pump absorbate contains the following components (the proportions are based on the total amount) Diphyl 52.55% by weight, dimethyl 12.11% by weight, diacrylate 1.38% by weight, acrylic acid 30.1% by weight, phenothiazine 0.03 wt.%, phthalic anhydride 0.08% by weight, benzoic acid 1.22% by weight, maleic anhydride 1.06 wt.%, benzaldehyde 0.61 wt.%, formic acid 0.03 wt.%, acrolein 0.007% by weight, water 0.20% by weight, acetic acid 0.28% by weight, propane 0.19% by weight propionic 0.01% by weight, formaldehyde 0.001% by weight, allyl 0.005 wt.%, furfural 0.02 wt .-% and methacrylic acid 0.07% by weight.

11.4 % By weight based on the total amount of absorbate becomes a desorption unit supplied, 3.1 Wt .-% based on the total amount directly to the bottom fed to the absorption column and 85.5 wt .-% are about series connected heat exchangers, in which it cooled to 15K is counted on the sixth bottom from below in the absorption unit recycled.

The in the first as tube bundle heat exchanger executed heat exchangers dissipated Heat becomes Partial evaporation of the condensation unit to be described later used accumulating sour water. The outlet temperature of the this heat exchanger leaving Absorbats 113.3 ° C. The further cooling of the absorbate to the inlet temperature of the return takes place in air coolers.

The Floors between the first chimney floor up to this inflow floor are as 4-flooded valve floors (steel VV12 without cover). The distance between two floors is 0.7 m. The floors above this inlet bottom to the 13th bottom are counted from below as 2-valve valve bottoms (Steel VV12 without cover) and are mounted at a distance of 0.5 m.

Above of the 13th floor counted from below is another designed as a double-walled fireplace ground catch bottom attached and over a centrifugal pump becomes a liquid one Withdrawable stream taken at a temperature of 71.8 ° C. 9 wt .-% of the The total quantity taken from this point will be immediately below this second catch bottom as return abandoned, wherein 1 wt .-% to 10 wt .-% based on this reflux amount be used to spray the second catch bottom via nozzles. The Remaining withdrawn amount is with the help of air coolers 59.4 ° C chilled and abandoned on the 19th bottom counted from below. The floors between the second catch bottom to the 19th bottom are counted from below as 4-valve valve shelves (steel VV12 without cover), where the distance between the floors is 0.6 m is.

In the withdrawal stream of the second catch bottom is also the on the Head of the rectification unit to be described below taken from acrylic acid Low-energy stream fed. Furthermore, this site is suitable in large-scale Method also for feeding acrylic acid-containing material streams, such as for example, unspecified crude acrylic acid or acrylic acid-containing Material flows from others Process stages, such as from the pure acrylic acid by distillation or crystallization.

The Floors above of the 19th floor counted from below up to the feed point of the fresh or slightly loaded absorbent in the absorption column are as 2-flow valve plates (steel VV12 without cover) with a distance between the floors of each executed 0.5 m. Above the 34th floor counted from below is a third as a double-walled fireplace floor executed catch bottom.

The remaining, not absorbed in the absorbent residual product gas mixture, that contains mainly low-boiling and non-condensable substances cooled in a condensation unit, which is above the third Fangboden joins. The condensation unit is designed as direct cooling and contains ten 2-valve valve bottoms (Steel VV12 without cover), spaced by 0.5 m exhibit.

For direct cooling already condensed aqueous low boiler fraction is withdrawn at the third catch bottom. 2% by weight, based on the total amount of acid water condensate withdrawn at the third collecting bottom, are discharged at a temperature of 45 ° C. and fed to an extraction of acid water. The discharged sour water has essentially the following contents: acrylic acid 4.02% by weight, acetic acid 5.04% by weight, water 69.84% by weight, Diphyl 12.95% by weight, dimethyl 2.51% by weight, formaldehyde 0.02 wt.%, diacrylate 0.2% by weight, allyl 0.004 wt.%, furfural 0.02 wt.%, propionic 0.0015% by weight, formic acid 0.62% by weight, benzaldehyde 0.84% by weight, maleic anhydride 0.05% by weight, benzoic acid 0.25% by weight, maleic 1.40% by weight, phthalic anhydride 0.01% by weight, phenothiazine 0.007% by weight, acrolein 0.02% by weight., methacrylic acid 0.05% by weight, methacrolein 0.005% by weight, Butane 0.21% by weight and butenes 0.08% by weight.

The Not discharged sour water condensate is in a tube bundle heat exchanger cooled to 33 ° C. 79 Wt .-% based on the amount of cooled to 33 ° C acid water condensate are abandoned on the 40th floor counted from below. The remaining one Quantity of sour water condensate is in another heat exchanger to 16 ° C cooled. This heat exchanger is energetic with the for the evaporation of the liquid Coupled propane necessary evaporator unit.

The in the acid water condensation gaseous residual gas mixture leaves the absorption unit at a temperature of 29 ° C and contains substantially following components:

nitrogen 50.40 Wt .-%, oxygen 3.57 Wt%., carbon oxides 3.65 % By weight, water 1.46 Wt%., acetic acid 0.058 Wt .-%, acrylic acid 0,041 Wt .-%, acrolein 0.067 Wt .-%, Diphyl 0,008 Wt .-%, formic acid 0.003 Wt .-%, Ethan 0.20 % By weight, ethene 0.03 % By weight, butenes 0.15 % By weight, Butane 0.55 % By weight, propane 39.62 Wt .-% and propene 0.17 Wt .-%.

Of the Remaining gas flow remaining in the acid water condensation (cycle gas) will over passed a demister from the column and in a tube bundle heat exchanger overheated by 6K. This will be a possible Condensation in the exhaust gas lines avoided. This residual gas stream is powered by an electrically driven turbo compressor on one Compressed pressure of 3.3 bar.

From This residual gas stream 20 vol .-% based on the total amount the residual gas stream as stripping gas in a later to be described Washing column used. The remaining residual gas flow is in the Reaction zone A for heterogeneously catalyzed propane dehydrogenation recycled.

desorption

The absorbate effectively discharged from the bottom most catch tray of the absorption unit becomes one Desorption supplied to free it in this still contained therein nen low-boiling components. The absorbate is first heated in a steam-heated shell and tube heat exchanger to 130 ° C and then fed to a desorption column at the top of the same. In this case, sprayed via nozzles 5 wt .-% absorbate, based on the total amount of the desorption column supplied absorbate, used to moisten the column walls, the manhole and the column hood.

When Column internals are used in the desorption column 38 Regensiebböden on which breakwaters are mounted. The hole diameter is the floors Counted 30 mm from bottom 1 from below to the bottom 9, 25 mm at the bottom 10 and 20 mm from bottom 11 to Head of the column. The ground clearance between the individual floors is 0.5 m. The top pressure of the desorption column is 1.83 bar.

The heat in the bottom of the column via two parallel operated evaporators with forced circulation.

When Stripping the desorption column is a partial stream of the residual gas stream from the head of the absorption unit, after being described in the later Washing column was washed, fed directly to the bottom. The Stripping gas is in this case before entering the desorption with a partial stream of the withdrawn at the bottom of the desorption and heated in the evaporator bottoms liquid mixed and in a tube bundle heat exchanger to 120 ° C preheated. The partial flow of the bottoms liquid which is mixed with the stripping gas is 9 wt .-% based on the Total amount of withdrawn bottoms liquid. Optionally, on this mixing with heated bottoms liquid also be waived.

The acrylic acid-laden absorbent obtained in the column bottom of the desorption column contains: Diphyl 62.8% by weight, dimethyl 14.7% by weight, diacrylate 1.8% by weight, acrylic acid 17.1% by weight, acetic acid 0.04 wt.%, formic acid 0.004 wt.%, propionic 0.006% by weight, phenothiazine 0.04 wt.%, phthalic anhydride 0.1% by weight, benzoic acid 1.5% by weight, maleic anhydride 1.1% by weight, methacrylic acid 0.06% by weight, benzaldehyde 0.62 wt .-% and furfural 0.02% by weight.

The Forced circulation evaporators are taken from the bottom of the column Bottoms liquid with a temperature of 131.5 ° C supplied and at 150 ° C heated.

One Partial stream of 63 wt .-% based on the total amount of withdrawn bottoms liquid is returned to the desorption column on the 8th bottom counted from below.

The stripping gas charged at the top of the desorption column and charged with low-boiling components is returned to the direct cooler used for cooling the product gas mixture B. The recirculated laden stripping gas also contains the fluiditribidity of the topmost bottom of the desorption column and contains essentially the following constituents: nitrogen 26.75% by weight, oxygen 1.89% by weight, carbon oxides 1.94% by weight, acrylic acid 36.84% by weight, Diphyl 7.57% by weight, water 0.95% by weight, acetic acid 0.56% by weight, dimethyl 1.39% by weight, formaldehyde 0.0% by weight, acrolein 0.01% by weight, formic acid 0.05% by weight, propionic 0.01% by weight, allyl 0.01% by weight, diacrylate 0.16 wt.%, phthalic anhydride 0.0% by weight, benzoic acid 0.15% by weight, maleic anhydride 0.47% by weight, benzaldehyde 0.27% by weight, Furturale 0.01% by weight, Ethan 0.1% by weight propane 20.4% by weight, propene 0.1% by weight butenes 0.06% by weight and Butane 0.23% by weight.

rectification

Of the taken from the desorption after the forced circulation evaporators Partial flow, the main from acrylic acid, Diphyl and dimethyl phthalate is added to a and a stripping section comprehensive rectification unit out.

The Rectification unit consists of 43 rain screen bottoms on which breakwater are mounted, the floors have different hole diameters: 50 mm from the 1st floor of counted down up to the 10th floor, 25 mm from the 11th floor to the 13th floor and 14th floor mm from the 14th floor to the top floor. The distance between the individual soils is each 0.4 m. The distance between 8th and 9th floor counted from the bottom is 1 m.

The Column is operated with a top pressure of 106 mbar. The feed of the loaded, partly of low-boiling components in the desorption unit liberated absorbent takes place via a loop with several Neck on the 8th bottom counted from below.

Below the 1st floor 430 Nm ³ / h air is supplied via a loop with multiple feed points of the column.

The heat supply in the bottom of the column via two parallel operated external circulation evaporator with forced circulation. The bottom temperature of the rectification unit is typically 188 ° C. The high boiler fraction condensing in the bottom of the rectification column contains essentially the following constituents: Diphyl 75.6% by weight, dimethyl 17.7% by weight, diacrylate 1.85% by weight, acrylic acid 0.79% by weight, phenothiazine 0.05% by weight, phthalic anhydride 0.1% by weight, benzoic acid 1.75% by weight, maleic anhydride 1.3% by weight, methacrylic acid 0.06% by weight, benzaldehyde 0.74 wt .-% and furfural 0.015% by weight.

The taken from the rectification unit, the high-boiling absorbent coded containing bottoms liquid becomes 83% by weight, referred to the feed to the rectification unit, discharged, and partly over a heat exchanger returned to the bottom region of the rectification column. The discharged high boiler fraction is passed through a solids separator (Cyclone) and optionally by fresh absorbent (Diphyl and dimethyl phthalate), the washing unit supplied. A small partial flow of 1.2% by weight based on the total amount the discharged bottoms liquid is fed to the sump area of the absorption.

27 plates above the feed into the rectification column is taken off via side draw a crude acrylic acid. The removal of the crude acrylic acid via a Regensiebboden with integrated Mittelabzugstasse. The withdrawn crude acrylic acid contains essentially the following constituents: acrylic acid 99.66 wt.%, acetic acid 0.135% by weight, water 0.007% by weight, formic acid 0.002% by weight, propionic 0.036 wt.%, furfural 0.022 wt.%, Allyacrylat 0.011% by weight, benzaldehyde 0.008% by weight, maleic anhydride 0.008% by weight, methacrylic acid 0.06% by weight, diacrylate 0.02 wt .-% and phenothiazine 0.025 wt .-%.

The withdrawn crude acrylic acid is cooled by means of two series-connected heat exchanger to 25 ° C. From the withdrawn crude acrylic acid be 15.5 wt .-% based on the feed to the rectification column discharged and a smaller partial flow is used as a solvent for the Polymerization inhibitor used.

The cooling of the light-ends stream separated off at the top of the rectification column takes place in two stages through two direct coolers. Both stages are designed as a DC quench, with the first stage condensate having a temperature of 52 ° C and the second stage condensate having a temperature of 24.5 ° C. The vapor line of the second condensation stage leads to the vacuum unit. In this case, two liquid ring pumps are operated in parallel, wherein according to DE-A-10143565 as a barrier liquid, the condensate from the second condensation stage is used. In a subsequent container, the liquid phase, primarily the ring liquid, is separated from the non-condensable portions which form the exhaust of the rectification unit. This has essentially the following composition: nitrogen 48% by weight, oxygen 14.5% by weight, acrylic acid 2.3% by weight, acetic acid 0.09 wt.%, water 0.25% by weight, acrolein 0.001% by weight, formic acid 0.04 wt.%, allyl 0,004Gew .-%, carbon oxides 0.03 wt.%, propane 33.65% by weight, propene 0.03% by weight, Butane 0.85% by weight and butenes 0.21% by weight.

The Exhaust gas from the rectification unit is burned together with other production residues. at high propane content of this vacuum exhaust gas, however, it is also conceivable attributed this current to the absorption.

The low boiler stream separated off at the top of the rectification column contains essentially the following components: acrylic acid 98.44% by weight, acetic acid 0.94% by weight, water 0.35% by weight, phenothiazine 0.028% by weight, diacrylate 0.027 wt.%, allyl 0.05% by weight, furfural 0.006% by weight, propionic 0.034 wt.%, methacrylic acid 0.01% by weight, formic acid 0.07 wt .-% and acrolein 0.001% by weight.

One Partial stream of the withdrawn as low-boiling fraction liquid of 47.6 wt .-% based on the feed to the rectification column is called return used that over several nozzles is abandoned. A small part is also used to by means of nozzles the column hood and the vapor line spray with inhibitor-containing condensate. 1.3% by weight of the low boiler fraction, based on the feed to the rectification column, is discharged and returned to the absorption unit as described therein.

When Polymerization inhibitor phenothiazine is used. The inhibitor is considered a solution with an inhibitor concentration of 1.1% by weight in the over sidestream prepared as described separated crude acrylic acid and the condensed Low boiler return and the condensates of the two direct cooling stages of the rectification unit continuously added. The solid stabilizer is in the form of Dandruff or pellets over screw conveyors in the stabilizer batch tank metered.

washing unit

The Washing column is operated at a top pressure of 3 bar. As a release Internals become rain screen floors used with a hole diameter of 30 mm. 30 of these soils are mounted with a ground clearance of 0.4 m.

When washing liquid becomes the bottoms liquid used from the rectification unit. The washing liquid is abandoned at the top of the wash column at a temperature of 50 ° C. Below the bottom bottom, this is the absorption unit incoming circulating gas is fed after it has previously reached the inlet temperature from 50 ° C chilled has been. The supplied Circulating gas amount is 37% by weight of the amount of washing liquid supplied.

The liquid obtained in the bottom of the column contains essentially the following components: Diphyl 75% by weight, dimethyl 17.6% by weight, diacrylate 1.9% by weight, acrylic acid 0.8% by weight, maleic anhydride 1.3% by weight, phenothiazine 0.05% by weight, phthalic anhydride 0.1% by weight, benzoic acid 1.7% by weight, acrolein 0.02 wt.%, water 0.17 wt.%, acetic acid 0.02 wt.%, benzaldehyde 0.74% by weight, furfural 0.015% by weight, methacrylic acid 0.06% by weight and propane 0.5% by weight.

The in the washing unit accumulating bottoms liquid is discharged and divided into two subsets in a weight ratio of 7.2 to 1. The larger partial flow is charged as described in the absorption column, while the small partial stream of acid water extraction is supplied.

The washed cycle gas essentially contains the following constituents: nitrogen 51.62% by weight, oxygen 3.66% by weight, carbon oxides 3.73% by weight, water 1.01 wt.%, acrylic acid 0.02 wt.%, Diphyl 0.03 wt.%, benzaldehyde 0.01% by weight, maleic anhydride 0.015% by weight, propane 38.96 wt.%, propene 0.16 wt.%, Butane 0.42% by weight, butenes 0.11% by weight and Ethan 0.2% by weight.

The washed circulating gas is used as stripping gas in the described desorption unit directed.

Acid water extraction

The Acid condensate from the absorption unit is in an extraction unit with her Part of the liquid drain extracted from the washing unit.

The Extraction unit consists of a stirred tank with a two-stage impeller and a horizontal settling tank. Inlet and outlet in settling tank are separated by internals transverse to the flow direction. The inlet temperature of the acid condensate in the extraction unit is 45 ° C. The mass ratio of Amount of bottom drain of the washing unit and sour water from the acid water condensation is 0.8 to 1.

The aqueous extract obtained in the extraction unit absorbs from the bottoms of the wash unit polar constituents such as diacrylic acid (Michael adduct) and maleic anhydride, the latter being hydrolyzed in the process. It essentially contains the following components: water 70-95% by weight, acetic acid 2-10% by weight, acrylic acid 1-5% by weight, formaldehyde 2-10% by weight, maleic 1-5% by weight, diacrylate 1-5% by weight, Diphyl 0.01-0.05 wt%, dimethyl 0.1-1% by weight, allyl 0.0005-0.002 wt%, furfural 0.001-0.001% by weight, formic acid 0.2-2% by weight, benzaldehyde 0.01-0.05 wt%, benzoic acid 0.05-0.2% by weight, maleic 0.5-3 wt .-% and phthalic anhydride 0.01-0.1% by weight.

The discharged aqueous Phase is burned along with the other production residues. In advance of his Combustion, the sour water is partially evaporated. The required Heat is transmitted through tube bundle heat exchangers taken from the absorption unit, where there on as described the first catch bottom condensate condensate cools.

The organic phase (the raffinate) obtained in the extraction unit contains essentially the following components: Diphyl 76.3% by weight, dimethyl 17.2% by weight, acrylic acid 1.1% by weight, maleic anhydride 0.35% by weight, diacrylate 0.85% by weight, phenothiazine 0.05% by weight, phthalic anhydride 0.1% by weight, benzoic acid 1.55% by weight, methacrylic acid 0.03 wt.%, propionic 0.02% by weight benzaldehyde 1.16% by weight, formic acid 0.05% by weight, acrolein 0.04 wt.%, water 0.7% by weight, acetic acid 0.41 wt .-% and furfural 0.02 wt .-%.

The discharged organic phase is in the absorption unit like described there fed.

High-boiling distillation

The described subset of the sump outlet of the absorption unit fed to a distillation unit and in this by heating separated into a high boiler fraction and a low boiler fraction. The distillation unit is designed in one stage and is at a pressure operated by 90 mbar. A technical embodiment in a distillation column but also known with separating agents known to those skilled in the art conceivable.

The heat over an outside one Forced circulation evaporator. Overheating the circulating liquid is there 5 K. In the distillation unit raises a temperature from 188 ° C one.

Out the circulation amount of the forced circulation evaporator becomes a subset taken and after dilution with a suitable diluent (For example, dimethylformamide or methanol) burned together with other production residues.

The high boiler fraction contains essentially the following components: Diphyl 25.8% by weight, dimethyl 43.2% by weight, phenothiazine 29.1% by weight, acrylic acid 0.13 wt.%, diacrylate 0.6% by weight, phthalic anhydride 0.3% by weight, benzoic acid 1% by weight, maleic anhydride 0.02 wt .-% and benzaldehyde 0.01% by weight.

From the high boiler fraction is 1.7% by weight, based on the distillation unit supplied feed quantity discharged and burned.

The in the distillation unit into the vapor phase converted low boiler fraction cooled by a combination of direct and indirect cooling and condensed. This is already condensed low boiler fraction in the vapor the distillation unit sprayed and together with the uncondensed vapor by means of a shell-and-tube heat exchanger chilled and the cooled ones liquid above the heat exchanger in the vapor sprayed. The cooler leaving stream has a temperature of 52 ° C and is in a container in gas and liquid phase separated.

The liquid low-boiling fraction obtained in the distillation unit contains essentially the following components: acrylic acid 4.5% by weight, Diphyl 54.8% by weight, dimethyl 36.7% by weight, diacrylate 1.25% by weight, acetic acid 0.04 wt.%, water 0.01% by weight, formic acid 0.005 wt.%, phenothiazine 0.165% by weight, benzaldehyde 0.12 wt.%, maleic anhydride 0.24% by weight, benzoic acid 1.89 wt.%, phthalic anhydride 0.14% by weight, furfural 0.003 wt .-% and acrolein 0.002 wt .-%.

Over a Centrifugal pump becomes the liquid Low boiler fraction as described above the shell and tube heat exchanger sprayed as described. 98.2 wt .-% of the withdrawn as low-boiling fraction liquid based on the fed to the distillation unit feed amount is discharged and returned to the absorption unit below the first chimney tray.

The exhaust gas leaving the distillation unit in gaseous state contains essentially the following components: acrylic acid 4.8% by weight, acetic acid 0.2% by weight, water 2.2% by weight, Diphyl 0.8% by weight, dimethyl 0.1% by weight, formaldehyde 0.13 wt.%, acrolein 0.13 wt.%, propionic 0.002% by weight, furfural 0.003 wt.%, benzaldehyde 0.07% by weight, maleic anhydride 0.1% by weight, benzoic acid 0.01% by weight, phthalic anhydride 0.01% by weight, diacrylate 0% by weight, formic acid 0.02 wt.%, allyl 0,003Gew .-%, Butane 1.1% by weight, butenes 0.3% by weight and propane 89.9% by weight.

This Exhaust gas is burned together with the other production residues.

The recovered crude acrylic acid can in a conventional manner in further process steps to a pure acrylic acid be further processed. Suitable for this purpose are a one-stage or, with a crude acrylic acid with significant proportions of low-boiling secondary components, a two-stage distillative process with prior aldehyde treatment by at least one primary Amino group-containing compound (such as hydrazine or aminoguanidine hydrogencarbonate) according to the teachings from EP-A 270 999, EP-A 648 732, as well as a crystallisative Method according to the teachings from EP-A 616 998, EP-A 792 867, EP-A 1 189 861 and WO 98/01404.

Becomes the pure acrylic acid in a crystallizing process step such as a layer crystallization won, so falls besides the pure acrylic acid also an acrylic acid Mother liquor / acid stream, as described above in the absorption unit for recovery of crude acrylic acid to be led back can.

example 1

It was proceeded as in Comparative Example 1. The (total) residual gas was not in the first heater upstream evaporator, but on the appropriate pressure compacted together with the 195 Nl / h of compressed air in the third Dehydrierreaktor upstream Heater led.

The content of CO ₂ of the reaction gas mixture leaving the third dehydrogenation reactor subsequently dropped from 3.25% by volume to 1.6% by volume.

This indicates that the procedure according to the invention is associated with a less than half reduced full combustion of the C ₃ hydrocarbons involved.

Further could this procedure according to the invention via a Period of 7500 hours of operation essentially without significant impairments operate.

Example 2

described here is the stationary Operating condition.

One designed as a horde reactor and adiabatically designed shaft furnace reactor points as the first section a reaction zone A two in the flow direction one behind the other arranged catalyst beds, each with a fixed bed a dehydrogenation catalyst according to Comparative Example 1 are loaded. Before each fixed bed is a static Gas mixer.

• a) rückgeführtes Produktgasgemisch A, das eine Temperatur von 617°C und einen Druck von 2,75 bar aufweist in einer Menge von 132872 Nm³/h, enthaltend 26,4 Vol.-% Propan, 6,6 Vol.-% Propen (Propylen), 0,1 Vol.-% H₂, 8,8 Vol.-% H₂O und 0 Vol.-% O₂ und
• b) 8758 Nm³/h Roh-Propan gemäß Vergleichsbeispiel 1.

The first catalyst bed in the flow direction, a reaction gas starting mixture is fed, which is composed as follows:

• a) recycled product gas mixture A, which has a temperature of 617 ° C and a pressure of 2.75 bar in an amount of 132872 Nm ³ / h, containing 26.4% by volume of propane, 6.6% by volume of propene (Propylene), 0.1 vol .-% H ₂ , 8.8 vol .-% H ₂ O and 0 vol .-% O ₂ and
• b) 8758 Nm ³ / h of crude propane according to Comparative Example 1.

The Temperature of raw propane is such that the temperature the starting reaction gas mixture is 560 ° C.

The reaction gas starting mixture amount is 141630 Nm ³ / h, with the following contents:
31.0% by volume of propane,
6.2% by volume of propene,
0.1% by volume H ₂ ,
8.2% by volume of H ₂ O and
0 vol.% O ₂

Of the Pressure of the starting reaction gas mixture is before entering the first Catalyst bed 2.75 bar.

The bed height of the first catalyst bed through which the reaction gas starting mixture flows is such that the reaction gas mixture leaves this fixed catalyst bed with the following contents:
26.4% by volume of propane,
9.5 vol.% Propene,
3.6% by volume H ₂ ,
7.9 vol .-% H ₂ O and
0 vol.% O ₂

The leaving amount is 146533 Nm ³ / h. The leaving temperature is 500 ° C and the leaving pressure is 2.65 bar. The in the first catalyst bed adjusting propane conversion is 11.8 mol% of the supplied propane.

6250 Nm ³ / h of air are metered into the reaction gas mixture leaving the first fixed catalyst bed as described. The air is preheated to 500 ° C and its pressure is such that the pressure of the reaction gas mixture resulting after the air dosing is 2.65 bar.

As it flows through the second fixed catalyst bed, first half of the molecular hydrogen contained in the reaction gas mixture is burned to water with the molecular oxygen added in the form of air. In this case, the reaction gas mixture is heated to 550 ° C. As the second catalyst fixed bed flows through, heterogeneously catalyzed dehydrogenation of the propane takes place. The bed height of the second fixed catalyst bed is dimensioned such that the reaction gas mixture leaves the second fixed catalyst bed with a temperature of 513 ° C. and a pressure of 2.55 bar with the following contents in an amount of 154746 Nm ³ / h:
22.9% by volume of propane,
11.9 vol.% Propene,
3.9 vol.% H ₂ ,
9.2% by volume of H ₂ O and
0 vol.% O ₂

761.6 Nm ³ / h of molecular hydrogen are first fed to this product gas mixture A *. Then 112792 Nm ³ / h of total residual gas from the separation zone B are fed at a temperature of 567 ° C. The total residual gas has the following contents:
31.1% by volume of propane,
0 vol.% Propene,
0% by volume H ₂ ,
1.9 vol .-% H ₂ O and
3.0 vol.% O ₂

The Supply of the total residual gas is carried out according to the principle of one with this Total residual gas as jet propelled jet pump, wherein the conveying direction through a motive nozzle over a Mixing section and a diffuser relaxed jet of propulsion in the second section of the reaction zone A and the suction of the Saugstutzens towards the mixture of molecular hydrogen and product gas mixture A * points.

The resulting reaction gas mixture A * flows in an amount of 267538 Nm ³ / h and at a temperature of 536 ° C and a pressure of 2.85 bar in a second shaft furnace reactor, which forms the second section of the reaction zone A, is also designed adiabatic and a fixed catalyst bed which is also charged with the dehydrogenation catalyst according to Comparative Example 1. It has the following contents:
26.3% by volume of propane,
6.4% by volume of propene,
2.6% by volume of H ₂ ,
6.1% by volume of H ₂ O and
1.3% by volume O ₂

At the Flow through This third fixed catalyst bed is first in the reaction gas mixture A * contained molecular oxygen with the molecular contained therein Hydrogen substantially completely burned to water, wherein the reaction gas mixture is heated to 620 ° C.

The Schütthöhe the third Fixed catalyst bed is such that in the reaction gas mixture heterogeneously catalyzed until leaving the same to a slight extent Propane dehydrogenation takes place.

As a result, 265745 Nm ³ / h of product gas mixture A with a temperature of 617 ° C. and a pressure of 2.75 bar leave the third fixed catalyst bed. It contains the following contents:
26.4% by volume of propane,
6.6% by volume of propene,
0.1% by volume H ₂ ,
8.8 vol .-% H ₂ O and
0 vol.% O ₂

The Product gas mixture A is in two halves of identical composition divided up.

From that will be one half as a component of the reaction gas starting mixture for the first Section of the reaction zone A in selbige recycled.

The other half will over an indirect heat exchanger W for gases the separation zone A supplied. in the heat exchangers becomes the recycled to the reaction zone A residual gas from the separation zone B to the 567 ° C heated.

In the separation zone B is from the second half of the product gas mixture A contained in the same propane and propene as in Comparative Example 1 absorbively separated and as in Comparative Example 1 in the reaction zone B of the two - stage heterogeneously catalyzed partial oxidation of Propens to acrylic acid subjected.

Out the resulting product gas mixture B, the acrylic acid in the Separation zone B as in Comparative Example 1 separated, and this remaining total residual gas compressed (to about 4 bar) and as described on the heat exchangers W supplied with the total residual gas as jet fuel jet pump and with the molecular hydrogen supplemented product gas mixture A * united. The amount of acrylic acid separated in the separation zone B is to 332 kmol / h. The amount of residual gas in the above-described Gas mixtures essentially consists of molecular nitrogen and from small amounts of carbon oxides. The method described is essentially long-lasting operable without reduction.

Example 3

described here is the stationary Operating condition.

One designed as a horde reactor and adiabatically designed shaft furnace reactor points as the first section a reaction zone A two in the flow direction one behind the other arranged catalyst beds, each with a fixed bed a dehydrogenation catalyst according to Comparative Example 1 are loaded. Before each fixed bed is a static Gas mixer.

• a) rückgeführtes Produktgasgemisch A, das eine Temperatur von 676°C und einen Druck von 1,75 bar aufweist in einer Menge von 106510 Nm³/h, enthaltend 12,3 Vol.-% Propan, 8,2 Vol.-% Propen (Propylen), 2,2 Vol.-% H₂, 7,7 Vol.-% H₂O und 0 Vol.-% O₂ und
• b) 8758 Nm³/h Roh-Propan gemäß Vergleichsbeispiel 1.

The first catalyst bed in the flow direction, a reaction gas starting mixture is fed, which is composed as follows:

• a) recycled product gas mixture A, which has a temperature of 676 ° C and a pressure of 1.75 bar in an amount of 106510 Nm ³ / h, containing 12.3% by volume of propane, 8.2% by volume of propene (Propylene), 2.2 vol .-% H ₂ , 7.7 vol .-% H ₂ O and 0 vol .-% O ₂ and
• b) 8758 Nm ³ / h of crude propane according to Comparative Example 1.

The Temperature of raw propane is such that the temperature of the reaction starting mixture 652 ° C is.

The reaction gas starting mixture amount is 115268 Nm ³ / h, with the following contents:
19.0% by volume of propane,
7.6% by volume of propene,
2.0% by volume H ₂ ,
7.1% by volume H ₂ O and
0 vol.% O ₂

Of the Pressure of the starting reaction gas mixture is before entering the first Catalyst bed 1.75 bar.

The bed height of the first catalyst bed through which the reaction gas starting mixture flows is such that the reaction gas mixture leaves this fixed catalyst bed with the following contents:
15.1% by volume of propane,
10.6% by volume of propene,
5.1% by volume H ₂ ,
6.9 vol .-% H ₂ O and
0 vol.% O ₂

The leaving quantity is 119308 Nm ³ / h. The leaving temperature is 590 ° C and the leaving pressure is 1.65 bar. The in the first catalyst bed adjusting propane conversion is 17.5 mol% of the supplied propane.

The reaction gas mixture leaving the first fixed catalyst bed as described is metered with 2598 Nm ³ / h of air. The air is preheated to 590 ° C and its pressure is such that the pressure of the reaction gas mixture resulting after the air dosing is 1.65 bar.

As it flows through the second fixed catalyst bed about 18% of the molecular hydrogen contained in the reaction gas mixture are first burned to water with the metered in the form of air molecular oxygen. The reaction gas mixture heats up to 618 ° C. As the second catalyst fixed bed flows through, heterogeneously catalyzed dehydrogenation of the propane takes place. The bed height of the second fixed catalyst bed is dimensioned such that the reaction gas mixture leaving the second fixed catalyst bed with a temperature of 546 ° C and a pressure of 1.55 bar with the following contents in an amount of 124387 Nm ³ / h:
12.0% by volume of propane,
12.6% by volume of propene,
6.6 vol.% H ₂ ,
7.5 vol .-% H ₂ O and
0 vol.% O ₂

89715 Nm ³ / h of total residual gas from the separation zone B are then fed to this product gas mixture A * at a temperature of 627 ° C. The total residual gas has the following contents:
14.6% by volume of propane,
0 vol.% Propene,
0% by volume H ₂ ,
1.9 vol .-% H ₂ O and
3.0 vol.% O ₂

The Supply of the total residual gas is carried out according to the principle of one with this Total residual gas as jet propelled jet pump, wherein the conveying direction through a motive nozzle over a Mixing section and a diffuser relaxed jet of propulsion in the second section of the reaction zone A and the suction of the Saugstutzens towards the mixture of molecular hydrogen and product gas mixture A * points.

The resulting reaction gas mixture A * flows in an amount of 214102 Nm ³ / h and with a temperature of 580 ° C and a pressure of 1.85 bar in a second shaft furnace reactor, which forms the second section of the reaction zone A, is also designed adiabatic and a fixed catalyst bed which is also charged with the dehydrogenation catalyst according to Comparative Example 1. It has the following contents:
13.1% by volume of propane,
7.3% by volume of propene,
3.8% by volume H ₂ ,
5.2% by volume of H ₂ O and
1.3% by volume O ₂

At the Flow through This third fixed catalyst bed is first in the reaction gas mixture A * contained molecular oxygen with the molecular glands contained therein Hydrogen substantially completely burned to water, wherein the reaction gas mixture heated to 692 ° C.

The Schütthöhe the third Fixed catalyst bed is such that in the reaction gas mixture heterogeneously catalyzed until leaving the same to a slight extent Propane dehydrogenation takes place.

As a result, 213021 Nm ³ / h of product gas mixture A with a temperature of 677 ° C. and a pressure of 1.75 bar leave the third fixed catalyst bed. It contains the following contents:
12.3% by volume of propane,
8.2% by volume of propene,
2.2% by volume of H ₂ ,
7.7 vol .-% H ₂ O and
0 vol.% O ₂

The Product gas mixture A is in two halves of identical composition divided up.

From that will be one half as a component of the reaction gas starting mixture for the first Section of the reaction zone A in selbige recycled.

The other half will over an indirect heat exchanger W for gases the separation zone A supplied. in the heat exchangers becomes the recycled to the reaction zone A residual gas from the separation zone B to the 627 ° C heated.

In the separation zone B is from the second half of the product gas mixture A contained in the same propane and propene as in Comparative Example 1 absorbively separated and as in Comparative Example 1 in the reaction zone B of the two - stage heterogeneously catalyzed partial oxidation of Propens to acrylic acid subjected.

Out the resulting product gas mixture B, the acrylic acid in the Separation zone B as in Comparative Example 1 separated, and this remaining total residual gas compressed (to about 4 bar) and as described on the heat exchangers W supplied with the total residual gas as jet fuel jet pump and combined with the product gas mixture A *. The separated in the separation zone B Amount of acrylic acid amounts at 332 kmol / h. The amount of residual gas in the above-described Gas mixtures essentially consists of molecular nitrogen and from small amounts of carbon oxides. The method described is essentially long-lasting operable without reduction.

US Provisional Patent Application No. 60 / 584,469, filed Jul. 1, 2004, is inserted in the present application by reference. In terms of There are many changes and deviations from the above teachings of the present invention possible. It can therefore be assumed that the invention, in the context of while adding to it Claims, other than specifically described herein.

Claims (41)

Process for the preparation of acrolein, or acrylic acid, or their mixture of propane, in which A) a first reaction zone A feeds at least two gaseous, propane-containing feed streams, of which at least one contains fresh propane, and in the reaction zone A her thus supplied propane a heterogeneously catalyzed dehydrogenation to give a propane and propylene-containing product gas mixture A is obtained, B) leads out of the reaction zone A product gas mixture A and separated in a first separation zone A from the components contained in the product gas mixture A, propane and propylene different components at least a subset and thereby remaining, propane and propylene containing, product gas mixture A 'C) used in a second reaction zone B for charging at least one oxidation reactor and in the at least one oxidation reactor contained in the product gas mixture A' propylene of a heterogeneously catalyzed gas phase partial oxidation with m O) to form an acrolein, or acrylic acid, or mixtures thereof as the target product and excess molecular oxygen containing product gas mixture B, D) from the reaction zone B product gas mixture B leads and separates in a second separation zone B contained in the product gas mixture B target product and of the remaining, unreacted propane, molecular oxygen and optionally unreacted propylene-containing residual gas at least one unreacted propane, molecular oxygen and optionally unreacted propylene containing subset as one of the at least two propane-containing feed streams in the reaction zone A, characterized in that this feedback in the Reaction zone A along the reaction path of the heterogeneously catalyzed dehydrogenation of the propane in the reaction zone A is carried out so that at the feed point already at least 5 mol% of the reaction zone A on the other Z. Feed streams fed propane are reacted in the reaction zone A dehydrating. Method according to claim 1, characterized in that that at least half of unreacted propane remaining in separation zone B, Molecular oxygen and optionally unreacted propylene-containing residual gas is recycled to the reaction zone A. Method according to claim 1, characterized in that that the total amount of remaining in the separation zone B, not converted propane, molecular oxygen and optionally unreacted propylene-containing residual gas in the reaction zone A is returned. Method according to one of claims 1 to 3, characterized that the repatriation of in the separation zone B remaining, unreacted propane, molecular Oxygen and optionally unreacted propylene containing Residual gas in the reaction zone A only at one point of the reaction zone A takes place. Method according to one of claims 1 to 3, characterized that the repatriation of in the separation zone B remaining, unreacted propane, molecular Oxygen and optionally unreacted propylene containing Residual gas in the reaction zone A via several successively arranged feed points distributed. Method according to one of claims 1 to 5, characterized that the repatriation of in the separation zone B remaining, unreacted propane, molecular Oxygen and optionally unreacted propylene containing Residual gas in the reaction zone A is carried out so that at the feed point at least 20 mol% of the propane fed to the reaction zone A via the other feed streams are reacted in the reaction zone A dehydrating. Method according to one of claims 1 to 5, characterized that the repatriation of in the separation zone B remaining, unreacted propane, molecular Oxygen and optionally unreacted propylene containing Residual gas in the reaction zone A is carried out so that at the feed point at least 30 mol% of the propane fed to the reaction zone A via the other feed streams are reacted in the reaction zone A dehydrating. Method according to one of claims 1 to 7, characterized the content of the recycled to the reaction zone A, remaining in the separation zone B, unreacted propane, molecular oxygen, and optionally unreacted propylene containing, residual gas at molecular Oxygen is 0.5 to 10 vol .-%. Method according to one of claims 1 to 7, characterized that the content of the recirculated to the reaction zone A, in the separation zone B Blei remaining, unreacted propane, molecular oxygen, and optionally unreacted propylene containing, residual gas at molecular Oxygen is 2 to 5 vol .-%. Method according to one of claims 1 to 9, characterized that the ratio the amount of propane, that of the reaction zone A via coming from the separation zone B. recycled residual gas is fed to the amount of propane, the reaction zone A over other supply streams in total supplied is 0.1 to 10. Method according to one of claims 1 to 10, characterized that the reaction zone A as propane-containing feed streams only from the separation zone B originating residual gas and crude propane as fresh propane supplied become. Method according to one of claims 1 to 11, characterized that in the reaction zone A reaction gas mixture at the feed station for the originating from the separation zone B residual gas, based on the in him contained molar amount of propane, a higher molar hydrogen content has, as the reaction zone A supplied reaction gas starting mixture. Method according to one of claims 1 to 12, characterized that is in the reaction zone A at the feed point for the the separation zone B derived residual gas from this supplied residual gas and before this residual gas supply at the feed point in the reaction zone A present reaction gas mixture forming reaction gas mixture, based on the contained in it molar amount of propane, a higher having molar hydrogen content than that of the reaction zone A. supplied Starting reaction gas mixture. Method according to one of claims 1 to 13, characterized the reaction temperatures in the reaction zone A are from 300 to 800 ° C. Method according to one of claims 1 to 14, characterized that in the reaction zone A dehydrogenation catalysts concomitantly used containing 10 to 99.9% by weight of zirconium dioxide, 0 to 60% by weight of aluminum oxide, Silica and / or titanium dioxide and 0.1 to 10 wt% at least an element of the first or second main group, an element the third subgroup, an element of the eighth subgroup of the periodic table of the elements, lanthanum and / or tin with the proviso that the sum % by weight gives 100% by weight. Method according to one of claims 1 to 15, characterized that in the reaction zone A as catalysts catalyst strands and / or Catalyst rings are used. Method according to one of claims 1 to 16, characterized in that a tray reactor is used as reaction zone A. Method according to one of claims 1 to 17, characterized that the conversion achieved in reaction zone A is based on propane on the reaction zone A total supplied propane and based on single pass through the reaction zone A, 30 to 60 mol% is. Method according to one of claims 1 to 17, characterized that the conversion achieved in reaction zone A is based on propane on the reaction zone A total supplied propane and based on single pass through the reaction zone A, 40 to 50 mol% is. Method according to one of claims 1 to 10 and 12 to 19, characterized in that the one formed in the reaction zone A. Product gas mixture A is divided into two subsets of identical composition, one of the two subsets as one of the propane-containing feed streams in the Reaction zone A returns and from the other subset of the reaction zone A out in the first separation zone A leads. Method according to claim 20, characterized in that that the recycled into the reaction zone A subset of the product gas mixture A, based on the total amount of product gas mixture formed in the reaction zone A. A, 30 to 70 vol .-% is. Method according to claim 20, characterized in that that the recycled into the reaction zone A subset of the product gas mixture A, based on the total amount of product gas mixture formed in the reaction zone A. A, 40 to 60 vol .-% is. Method according to one of claims 20 to 22, characterized that the recycled into the reaction zone A subset of the product gas mixture A as a component of the reaction zone A supplied starting reaction gas mixture is recycled to the reaction zone A. Method according to one of claims 1 to 23, characterized that one designs and operates the reaction zone A with the proviso that it consists of a first and a second section, wherein you I. the first section of the reaction zone A at least one, gaseous Propane-containing feed stream containing fresh propane, and in this first section of the reaction zone A, the propane so supplied a heterogeneously catalyzed dehydrogenation, wherein a Propane, propylene and molecular hydrogen containing product gas mixture A * is obtained, upon receipt of which at least 5 mol% of the first Section of the reaction zone A supplied propane in this first Section of the reaction zone A have been reacted dehydrating and that, based on the molar amount of propane contained in it, one larger molars Amount of molecular hydrogen contains, than that of the first section the reaction zone A supplied Starting reaction gas mixture; II. The product gas mixture A * then molecular Oxygen, unreacted propane and optionally unreacted Propylene-containing residual gas from the separation zone B feeds and the resulting reaction gas mixture A * the second section the reaction zone A supplies, and III. in this second section of the reaction zone A below Formation of the propane and propylene-containing product gas mixture A molecular oxygen contained in the reaction gas mixture A * with heterogeneous in the reaction gas mixture A * contained molecular hydrogen catalyses burned to water and contained in the reaction gas mixture A * Propane optionally dehydrogenated heterogeneously catalyzed to propylene. Method according to Claim 24, characterized that the first section of the reaction zone A is made adiabatic. Method according to claim 24 or 25, characterized that the second section of the reaction zone A adiabatic designed becomes. Method according to one of claims 24 to 26, characterized that both the first and the second section of the reaction zone A with the proviso adiabat is designed that - the one-time passage the reaction gas starting mixture fed to the first section of reaction zone A. gross calorific value obtained by the first section of reaction zone A is endothermic until autothermic, and - the on the one-time passage of the second section of the reaction zone A supplied Reaction gas mixture A * through the second section of the reaction zone A gross calorific value is exothermic. Method according to one of Claims 24 to 27, characterized the first section of the reaction zone has a horde structure. Method according to one of claims 24 to 28, characterized that the product gas mixture A less than 5% by volume of molecular hydrogen contains. Method according to one of Claims 24 to 29, characterized that the first and the second section of the reaction zone A in a Reactor are housed. Method according to one of claims 24 to 30, characterized that the product gas mixture A in two subsets of identical composition split and the one of the two subsets in the first section the reaction zone A is recycled. Method according to one of claims 24 to 31, characterized that the product gas mixture A *, the unreacted propane, molecular oxygen and optionally unreacted propylene Containing residual gas from the separation zone B according to the principle of this residual gas supplied as a jet propelled jet pump, wherein the conveying direction through a motive nozzle over a Mixing section and a diffuser relaxed jet of propulsion in the second section of the reaction zone A and the suction of the suction nozzle towards the first portion of the reaction zone A and the connection suction nozzle mixed-route diffuser the sole link forms between the two sections of the reaction zone A. Method according to claim 32, characterized in that that one chooses the pressure of the propulsion jet so that the pressure of the in the second section of the reaction zone A forming product gas mixture A above the pressure in the first section of the reaction zone A forming product gas mixture A * is. Method according to claim 33, characterized that the product gas mixture A in two subsets of identical composition split and one of the two subsets in a kind of natural circulation the pressure gradient between the product gas mixture A and the product gas mixture A * following is recycled to the first section of reaction zone A. Method according to one of claims 1 to 34, characterized that is guided into the separation zone A. Product gas mixture over an indirect heat exchanger led into the separation zone A. is going to be in this indirect heat exchanger to heat the fresh propane and / or the residual gas from the separation zone B. Method according to one of claims 1 to 35, characterized that neither of the two reaction zones A, B and neither Separation zones A, B external water vapor is supplied. Method according to one of claims 1 to 36, characterized that the product gas mixture A in the separation zone B with an organic solvent is contacted in which propane and propylene absorbs become. Method according to one of claims 1 to 37, characterized that the reaction zone B as a two-stage heterogeneously catalyzed Gas-phase partial oxidation from propylene to acrylic acid is designed. Method according to one of claims 1 to 38, characterized in that the Zielproduktabtren tion in the reaction zone B by absorption of the target product in an absorbent or by fractional condensation of the product gas mixture B. Method according to one of claims 1 to 39, characterized that remaining in the separation zone B and in the reaction zone A recycled residual gas having the following constituents with the following contents: 10 up to 40% by volume of propane, 0 to 1% by volume of propene, > 0 to 5 vol .-% molecular Oxygen, 1 to 10 vol .-% water vapor and 0 to 0.5 Vol .-% molecular hydrogen. Method according to one of claims 1 to 40, characterized that in the reaction zone A, the molar ratio of in the reaction gas mixture Propylene contained in the reaction gas mixture contained molecular Hydrogen does not exceed the value 10 at any point.