TWI508941B - Process for preparing acrolein or acrylic acid or a mixture thereof from propane - Google Patents

Description

Method for preparing acrolein or acrylic acid or a mixture thereof from propane

The present invention relates to a process for the preparation of acrolein or acrylic acid or a mixture thereof from propane, wherein A) - a gaseous feed stream comprising at least two propanes is sent to the first reaction zone A to form a reaction gas A, of which at least two At least one of the feed streams comprises fresh propane; - in reaction zone A, directing reaction gas A through at least one catalyst bed, wherein a portion of the propane partially catalyzes dehydrogenation to form molecular hydrogen and propylene; The molecular oxygen is sent to the reaction zone A, and at least a part of the molecules present in the reaction gas A are oxidized to steam in the reaction zone A, and the product gas A including propylene, propane and steam is extracted from the reaction zone A; Where appropriate, in the first separation zone I, the vapor present in the product gas A is partially or completely condensed by the product gas A and/or direct cooling to leave the product gas A ^* ; C) in a reaction zone B, a gaseous product a or product gas a ^* packed ^with molecular oxygen in the feed comprises at least one oxidation reactor having propane, propylene and molecular oxygen react the gas B, and wherein the propylene present in the The heterogeneously catalyzed partial gas phase oxidation to obtain acrolein or acrylic acid or a mixture thereof as a target product, and a product gas B including unconverted propane; D) deriving product gas B to reaction zone B, and in the second separation zone Part II removes the target product present therein to leave a residual gas comprising propane; E) if appropriate, a portion of the residual gas having a residual gas composition is recycled to reaction zone A as a feed stream comprising propane; In a separation zone III, the propane present in the residual gas is absorbed from the residual gas into the organic solvent (absorbent) by absorption to form an absorbent comprising propane, wherein the residual gas has not been recycled to the reaction zone A Any of the vapors present therein and where appropriate, have been previously partially or completely removed by condensation and/or any molecular hydrogen present therein has been partially or completely removed by means of a separation membrane in advance; and G) is separated In zone IV, propane is removed from the absorbent and recycled to reaction zone A as a feed stream comprising propane.

Acrylic acid is an important bulk chemical which is especially used as a monomer for preparing a polymer which is, for example, dispersed in an aqueous medium as a binder. Another field of use of acrylic polymers is in the field of use of superabsorbents in sanitary articles and in other fields of use. Acrolein is, for example, an important intermediate for the preparation of glutaraldehyde, methionine, 1,3-propanediol, 3-methylpyridine, folic acid and acrylic acid (see, for example, US-A 6,166,263 and US-A 6,187,963). .

A process for the preparation of acrolein or acrylic acid or a mixture thereof from propane is known in which, in the first reaction zone, propane is partially dehydrogenated to propylene under heterogeneous catalysis, and the propylene formed is subsequently In the presence of the converted propane (i.e., as recommended in, for example, EP-A 1 617 931 (see also WO 04/09041), there is no need to remove propylene from the propane remaining in the pre-dehydrogenation), in heterogeneous catalysis Partial oxidation under action to obtain acrolein or to obtain acrylic acid or to obtain a mixture thereof (see, for example, DE-A 33 13 573, EP-A 117 146, US-A 3,161,670, DE-A 10 2004 032 129, EP-A 731 077, DE-A 10 2005 049 699, DE-A 10 2005 052 923, WO 01/96271, WO 03/011804, WO 03/076370, WO 01/96270, DE-A 10 2005 009 891, DE-A 10 2005 013 039, DE-A 10 2005 022 798, DE-A 10 2005 009 885, DE-A 10 2005 010 111, DE-A 102 455 85, DE-A 103 160 39, DE 10 2005 056377.5, DE 10 2005 049699.7, DE 10 2006 015235.2, DE 10 2005 061626.7 and the prior art cited in these references).

A common feature of the method described is that after removing the target product acrolein or acrylic acid or acrolein and acrylic acid from the partially oxidized product gas mixture, at least a portion of the propane is present in the residual gas from the remaining residual gas including propane. The recycle gas is recycled to the heterogeneously catalyzed partial dehydrogenation of propane to propylene.

Additionally, the prior art methods described above can be substantially divided into two groups. The method of one of the two groups is characterized in that the components other than propane and propylene are removed by heterogeneously catalyzed partial dehydrogenation of the product gas mixture by absorption, and then the propylene is partially oxidized in the presence of residual propane. The desired target product (DE 10 2005 013039.9). A disadvantage of this procedure is that it typically requires at least 3 zones where the gas must be compressed. First, it generally requires compression of the heterogeneously catalyzed partial dehydrogenation of the product gas of propane, after which it is removed by the absorption of propylene and propane present therein, since absorption removal can generally be performed efficiently only under pressure. . In addition, it generally requires compression of the source of oxygen for the heterogeneously catalyzed partial oxidation, and in addition, it typically requires compression remaining in the heterogeneously catalyzed partial oxidation of propylene and recycled to the heterogeneously catalyzed partial dehydrogenation. Propane. The method of the other of the two groups is characterized in that it is recommended by those skilled in the art to use a heterogeneously catalyzed partial dehydrogenation product gas mixture as described to charge the heterogeneously catalyzed partial oxidation of propylene (see, for example) DE-A 103 160 39). The basis of the procedure (especially) is the assumption that molecular hydrogen formed in the heterogeneous partial catalytic dehydrogenation of propane does not have an adverse effect in the subsequent partial catalytic partial oxidation of propylene (and promoted by the presence of oxygen) and wherein the free hydrogen is not formed as an intermediate (to be withdrawn from the dehydrogenation of hydrocarbons to hydrogen water (H ₂ O) extracted directly) detectable nor to the uneven heating contrast phase catalytic oxydehydrogenation, uneven phase catalytic Dehydrogenation (e.g., the one to be performed in reaction zone A) will be understood in the present application to mean ("preferred") dehydrogenation, in contrast to oxidative dehydrogenation, which has an endothermic characteristic ( The expiratory hydrogen combustion may be included in the reaction zone A as a subsequent step) and wherein the free molecular hydrogen is formed at least as an intermediate; it generally requires different reaction conditions and different catalysts for oxidative dehydrogenation; in other words, in accordance with the present invention the method of propane in the reaction zone A portion of the uneven phase catalytic dehydrogenation must be precipitated under H _2). However, in-house studies have shown that the increased amount of molecular hydrogen in the presence of a reactive gas mixture for the heterogeneously catalyzed partial oxidation of propylene, while having no significant effect on the partial oxidation of by-product spectra, is indeed due to the molecular The significant reduction potential of hydrogen) significantly reduces the lifetime of the catalyst to be used for the heterogeneously catalyzed partial gas phase oxidation.

In addition, molecular hydrogen causes the reactive gas B to carry a potentially explosive component with a specific diffusion behavior (some materials of construction are permeable to molecular hydrogen). The use of excess molecular oxygen is inextricably linked to the above problems, which is advantageous for the increased catalyst life in the partial oxidation of reaction zone B relative to the target reaction stoichiometry, which of course brings about the presence of molecular hydrogen in Increased risk of explosion in the case of reaction zone B (when acrylic acid is the target product and the total amount of molecular oxygen required is added to the reaction gas B in advance, at least >1.5 of the molecular oxygen present in the reaction gas B is usually required The molar ratio to the propylene present therein (which has been considered to a low degree of complete propylene combustion)).

In view of the prior art described, it is an object of the present invention to provide an improved method having the disadvantages described in simplified form in the worst case according to the preamble herein.

Thus, a process for the preparation of acrolein or acrylic acid or mixtures thereof from propane has been found in which A) - a gaseous feed stream comprising at least two propanes is sent to a first reaction zone A to form a reaction gas A, at least two of which At least one of the streams comprises fresh propane; - in reaction zone A, directing reaction gas A through at least one catalyst bed, wherein a portion of the propane is heterogeneously catalytically dehydrogenated to form molecular hydrogen and propylene; The molecular oxygen is sent to the reaction zone A, and at least a part of the molecules present in the reaction gas A are oxidized to steam in the reaction zone A, and the product gas A including propylene, propane and steam is extracted from the reaction zone A; Where appropriate, in the first separation zone I, the vapor present in the product gas A is partially or completely condensed by the product gas A and/or direct cooling to leave the product gas A ^* ; C) in reaction zone B, a gaseous product a or product gas a ^* packed ^with molecular oxygen in the feed comprises at least one oxidation reactor having propane, propylene and molecular oxygen react the gas B, and propylene present in the warp wherein The phase-catalyzed partial gas phase oxidation to obtain acrolein or acrylic acid or a mixture thereof as a target product, and a product gas B including unconverted propane; D) deriving product gas B to reaction zone B and in second separation zone II Removing the target product present therein to leave a residual gas comprising propane; E) recycling a portion of the residual gas having a residual gas composition to the reaction zone A as a feed stream comprising propane, as appropriate (residual gas circulation) Gas); F) in the separation zone III, the propane present in the residual gas is absorbed from the residual gas into the organic solvent (absorbent) by absorption to form an absorbent comprising propane, wherein the residual gas has not been recycled Any molecular hydrogen in the reaction zone A, and optionally any steam present therein, which has been previously partially or completely removed by condensation and/or present therein, has been partially or completely removed by means of a separation membrane in advance; G) in the separation zone IV, the propane is removed from the absorbent and recycled to the reaction zone A as a feed stream comprising propane, preferably a propane recycle gas comprising steam; A in a reaction zone at least a sufficient molecular hydrogen to the steam in the reaction zone A as the amount of hydrogen peroxide in an amount of steam of formation of molecular hydrogen in the reaction zone A, at least 20 mol%.

Advantageously, according to the invention, at least some of the thermal energy generated by the oxidation (combustion) of molecular hydrogen in the reaction zone A is also used by heat exchange with the reaction gas A and/or the product gas A as a heat carrier, The component of the gas mixture is charged with the reaction gas A heated (sent to the gaseous feed stream of reaction zone A).

Advantageously, according to the present invention, the amount of hydrogen oxidized to vapor in the reaction zone A is at least 25 mol%, more preferably at least 30 mol%, based on the amount of molecular hydrogen formed in the reaction zone A under each condition. At least 35 mol%, more preferably at least 40 mol%, even more preferably at least 45 mol% and most preferably at least 50 mol%.

In other words, the method according to the invention is also a method wherein, in the reaction zone A, up to 55 mol%, or up to 60 mol%, or up to 65, based on the amount of molecular hydrogen formed in the reaction zone A Mol%, or up to 70 mol%, or up to 75 mol%, or up to 80 mol%, or up to 85 mol%, or up to 90 mol%, or up to 95 mol%, or up to 100 mol The molecule of % is oxidized to steam.

When the molecular hydrogen, which is a component of the reaction gas B, has an effect on the life of the catalyst in partial oxidation, the molecular hydrogen is combusted (oxidized) into water to obtain about twice the amount of hydrogen consumed in the dehydrogenation process of propane. In the case of heat, in terms of heat balance, when in the method according to the present invention, the amount of molecular hydrogen burned in the reaction zone A is 30 to 70 mol%, based on the amount of molecular hydrogen formed in the reaction zone A, It is advantageous when it is 40 to 60 mol%.

The limited content of molecular hydrogen in the reaction gas B (i.e., the reaction gas B which still includes molecular hydrogen) is also advantageous to some extent because molecular hydrogen has the advantage of the highest thermal conductivity among the gases. According to Walter J. Moore, Physikalische Chemie [Physical Chemistry], WDEG Verlag, Berlin (1973), p. 171, the thermal conductivity under standard conditions (for example) is greater than 10 times the thermal conductivity of carbon dioxide and molecular nitrogen or molecular oxygen. The thermal conductivity is about 8 times larger.

According to DE-A of Table I 33 13 573, the increase in thermal conductivity of the reaction zone B, is selectively formed in the presence of molecular hydrogen significantly lower than that of CO _x selectivity in the absence of molecular hydrogen to CO _x formation of the reason. Since the heat is removed from the reaction sites in the reaction zone B more quickly, it causes a lower catalyst surface temperature and thus a lower proportion of propylene is completely burned. Especially as in the method according to the invention, it is correct when the molecular hydrogen having the highest thermal conductivity and the propane having the highest heat absorption rate are synergistically included in the heat removal.

However, in principle, the reaction gas B in the process according to the invention may no longer comprise any molecular hydrogen.

In principle, product gas A comprising propylene, propane and steam and withdrawn from reaction zone A can be used as indicated to charge at least one oxidation reactor in reaction zone B. According to the invention, however, by direct and/or indirect cooling of the product gas A, at least a portion of the steam present in the product gas A in the first separation zone I is condensed, and thus from the product gas A It is removed to leave a product gas A ^* whose vapor content is lower than the vapor content of the product gas A.

One of the reasons for this is that the acrolein and/or acrylic acid target product has a significantly higher affinity for water than the component of product gas A, which is a significant removal of the target product from product gas B including high amounts of steam. The reasons related to energy demand. Secondly, the reaction volume in the gas steam typically reduces B generally comprises the oxidized form and Mo catalysts for the partial oxidation of life, since H ₂ O to promote the sublimation of the molybdenum oxides. Additionally, the vapor removed between reaction zone A and reaction zone B reduces the total volume of reactive gas B, and thus reduces the energy requirement to be delivered through reaction zone B.

Advantageously, at least 5 mol%, more preferably at least 10 mol%, even more preferably at least 15 mol%, especially advantageously at least 20 mol%, more advantageously at least 25 mol% or at least 30 mol%, present in the product gas A Even more advantageously at least 35 mol% or at least 40 mol%, more preferably at least 45 mol% or at least 50 mol%, particularly advantageously at least 55 mol% or at least 60 mol%, preferably at least 65 mol% or at least 70 Mol%, more preferably at least 75 mol% or at least 80 mol%, even more preferably at least 85 mol% or at least 90 mol%, in many cases at least 95 mol% or in extreme cases up to 100 mol% (but Frequently no more than 80 mol%, or no more than 85 mol%, or no more than 90 mol%, or no more than 95 mol%, or no more than 98 mol% of steam by means of the separation zone 1 in the process according to the invention Removed by condensation. The limited (but not zero) vapor content in the reactive gas B is advantageous because it generally has a beneficial effect on the activity of the catalyst used in the reaction zone B (especially in the case of multimetal oxides).

Advantageously, the removal of water between reaction zone A and reaction zone B can be achieved by simple condensation of the process according to the invention.

The water removed between the reaction zone A and the reaction zone B may be not only water formed by the oxidation in the reaction zone A. In fact, steam may also be additionally used as the inert diluent gas in reaction zone A and, if appropriate, also partially or completely removed in separation zone I.

Especially at high pressures, since the boiling point of water is quite high, the above condensation can usually be achieved by simple direct and/or indirect cooling of the product gas A. In principle, condensation can be achieved by single or multi-stage cooling. Suitably, condensation of water will be achieved by multi-stage cooling in accordance with the present invention. Advantageously from the application point of view, direct and indirect cooling will be used in combination. Suitable indirect heat exchangers for the purposes of the present invention are in principle all known types of indirect heat exchangers. Among these heat exchangers, a tube bundle heat exchanger and a plate heat exchanger are preferably used. For example, in the first cooling stage, the product gas A thermally extracted from the reaction zone A can be directed (the typical temperature of the product gas A is, for example, 400 to 700 ° C, preferably 500 to 600 ° C; typical of the product gas A The working pressure is usually >1 or 1.5 to 5 bar, or 2 to 5 bar, preferably 1.5 or 2 to 3 bar) in the first indirect heat exchanger (more generally, the hot product gas A withdrawn from the reaction zone A is available In the method according to the invention, at least some of the gas mixture is heated by means of indirect heat exchange, in a series of heat exchangers connected in series (for example, two heat exchangers connected in series), as appropriate Component or all components (to the desired temperature in each case) and simultaneously cooling the product gas A)), for example with a gas stream comprising propane and having been removed from the absorption in separation zone IV (propane recycle gas) Cocurrent or countercurrent to achieve the desired inlet temperature level of reaction zone A as one of the feed streams comprising propane introduced into reaction zone A. This temperature level can be typically used in a manner according to the method of the invention, for example, in the range of 400 to 600 ° C, preferably 450 to 550 ° C. In the downstream second indirect heat exchanger, the product gas A which is cooled as described will be directed, for example, in the same or countercurrent to the product gas A which has been previously removed to the desired extent by the steam present therein (in other words, It is directed to co-current or countercurrent to the product gas A ^* to raise it again to a suitable and suitable temperature level for subsequent partial oxidation of the propylene present therein. When it leaves the second indirect heat exchanger, the product gas A will still typically have a temperature above 100 °C. In a third indirect heat exchanger that is typically equipped as an air cooler (eg, product gas A is co-current or counter-current with air), the temperature of product gas A will then be lowered to a temperature below 100 °C. The removal of steam can then be carried out in a subsequent direct cooler. The direct cooler (direct condenser) may have a design known per se (for example, it may correspond to a rectification column) and may have conventional internal components for the purpose of enlarging the heat exchange surface area (eg, customary for the interior of the rectification column) member). However, in general, such internal components will be omitted. Typically, the tower of the direct condenser is not insulated. Preferred are all measures to increase the heat loss through the tower wall, such as perforated sheets or cooling ribs welded parallel to the surface of the tower. The coolant for the purpose of direct cooling of the product gas A will advantageously be an aqueous phase which has been previously removed from other product gases A, which is suitably cooled by means of cooling water before being used for direct cooling (from an application point of view) Look properly, surface water), cooled in an indirect heat exchanger to 35 ° C, preferably 30 ° C temperature. In a direct cooler, the product gas A and the finely atomized aqueous direct coolant can be directed in a cocurrent or countercurrent flow. In general, it is preferred to flow in the same stream.

The aqueous phase and the remaining gas phase can be separated from one another by means of a separator known per se (for example by means of a mechanical droplet separator). After the cooling of the removed aqueous phase is performed as described, the aqueous phase can be sent again to the direct cooling of the product gas A.

Remaining vapor (in the condition, which is the product gas A ^*) of the temperature may be raised to the subsequent adapted by indirect heat remains free of product vapors of the gas exchange A (described above) of propylene The temperature of subsequent partial oxidation. In principle, in the process according to the invention, the steam present in the product gas A may also be removed from only a portion thereof by condensation. The product gas A ^* thus formed can then be used, as indicated or in a mixture with a portion of the product gas A which has not been subjected to condensation treatment (in the sense of the invention, the mixture is likewise a product gas A ^* ) The heterogeneous phase of the loading of propylene catalyzes partial oxidation.

Another advantageous effect of direct cooling of product gas A (for example, by condensing the aqueous phase removed from product gas A, preferably by indirect heat exchange) is usually direct cooling while causing the presence of product gas. Deposition of any solid particles in A, which may be, for example, dehydrogenation catalyst particles derived from the heterogeneously catalyzed partial dehydrogenation of propane, which may be in the form of propylene according to the teachings of DE-A 103 160 39 Subsequent heterogeneously catalyzed partial oxidation has a cumbersome effect. It is especially true when the product gas A or the product gas A ^* used to charge the partial oxidation still includes molecular hydrogen. In other words, according to the invention, the mechanical separation operation according to DE-A 103 160 39 is more generally connected between the reaction zone A and the reaction zone B, even if only the direct cooling of the reaction gas A to be carried out as described is carried out. form.

In the separation zone I, the condensation of the vapor from the product gas A is removed, which comprises directly cooling the product gas A when it comprises using an aqueous phase which is preferably removed from the product gas A and cooled (preferably by indirect heat exchange). It is extremely especially advantageous.

The aqueous phase removed by condensation will typically still comprise dissolved propane, which may be recovered in a simple manner, for example by heating the aqueous phase (where appropriate under reduced pressure) and as another comprising propane The stream is recycled to reaction zone A. Direct and/or indirect cooling of product gas A is generally associated with the purpose of condensate removal of water present in product gas A. 1 bar, usually 0.5 bar, preferably The pressure drop of 0.3 bar is related.

Hydrogenation of the molecule into water in reaction zone A can be carried out after the actual heterogeneously catalyzed partial dehydrogenation of propane and/or during the heterogeneous partial catalytic dehydrogenation of propane, for example when in reaction gas A The conversion rate of propane produced in the reaction zone A in a single pass is performed when some partial conversion values have been obtained. Advantageously, according to the invention, at least a portion of the hydrogen oxidized to water in reaction zone A has obtained the desired target conversion of propane (and the resulting end point conversion) in reaction zone A (single reaction zone A is passed through reaction zone A) Oxidized into water before.

This procedure is advantageous first because the molecular hydrogen is withdrawn by equilibrium from the heterogeneously catalyzed dehydrogenation, which shifts the equilibrium position towards the desired propylene. However, it is also advantageous because the heat of reaction obtained in the oxidation again compensates for the heat of reaction consumed in the previously achieved dehydrogenation and provides the heat of reaction for subsequent further dehydrogenation. Advantageously, according to the present invention, molecular hydrogen (at least in part) will be formed at 90 mol%, more preferably 75 mol%, of the total amount of molecular hydrogen formed in the reaction zone A, and molecular oxygen in the reaction zone A. combustion. According to the present invention, it is preferred that the molecular hydrogen (at least partially) will be formed in the reaction zone no later than 65 mol%, or 55 mol%, or 45 mol% of the total amount of molecular hydrogen formed in the reaction zone A. The molecular oxygen in A burns.

The time during which the molecular hydrogen is burned in the reaction zone A can be first affected by the position (time) at which the molecular oxygen is sent to the reaction zone A.

When a portion of the residual gas including propane remaining in the separation zone II (the composition of which corresponds to the composition of the residual gas) is directed to (ie, does not pass through the separation zone III/IV) reaction zone A, the reaction zone A The filling gas mixture may already include molecular oxygen since the residual gas described above typically still includes molecular oxygen that is used in excess in reaction zone B.

Alternatively, it may be affected by the configuration in which the catalyst zone A is loaded. Thus, reaction zone A can be partially loaded with a catalyst which only specifically catalyzes the oxidation of molecular hydrogen and molecular oxygen to water. Other catalysts are capable of catalytically dehydrogenating selectively. In general, such catalysts capable of catalyzing dehydrogenation are also capable of catalyzing the combustion of hydrogen, in which case it is generally possible to reduce the activation energy of the latter reaction, in particular significantly.

In the present context, fresh propane is understood to mean propane which has not yet participated in the dehydrogenation in reaction zone A. In general, it is supplied as a component which also contains a small amount of crude propane in addition to propane (which preferably satisfies the specifications of DE-A 10246119 and DE-A 10245585).

The crude propane can be obtained, for example, by the method described in DE-A 102005 022798. In general, in addition to fresh propane, the feed stream which additionally feeds to the reaction zone A in the process according to the invention, which additionally comprises propane, is the propane removed from the absorption in the separation zone IV (and, where appropriate, the residual gas) One part of the circulating gas).

In the process according to the invention, it is preferred to feed fresh propane as a component of the filling gas mixture of reaction zone A to reaction zone A. In principle, part of the fresh propane may also be sent to the first and/or second oxidation stage of reaction zone B or to the filling gas mixture at any other point of the process for reasons of explosion safety.

In this context, the loading of the catalyst bed to the reaction gas in the catalytic reaction step is understood to mean a standard liter (= 1 (STP)) per hour through a 1 liter catalyst bed (eg fixed catalyst bed); The amount of the reaction gas in the lower (0 ° C, 1 bar), the corresponding amount of the reaction gas in the volume of the liters.

The loading may also be based on only one component of the reaction gas. In this case, it is 1 (STP)/1 per hour through a 1 liter catalyst bed. h The amount of this component (the fixed catalyst bed does not take into account the bed of pure inert material). When the catalyst bed consists of a mixture of a catalyst and an inert shaped diluent body, when the loading is referred to herein, it may also be based only on the volume unit of catalyst present.

As used herein, inert gas is generally understood to mean a reactive gas component which, under the respective reaction conditions, is substantially chemically inert and remains chemically unaltered to greater than 95 upon individual consideration of each inert reactive gas component. The degree of mol% is preferably greater than 97 mol% or greater than 99 mol%.

In addition to the residual gas recycle gas from separation zone II, the suitable source of molecular oxygen required in reaction zone A is a molecular oxygen or molecular oxygen as indicated in the reaction zone A (or chemically inert in reaction zone A) (or A mixture of such inert gases (for example, a rare gas such as argon, molecular nitrogen, steam, carbon dioxide, etc.). In other words, the molecular oxygen may be sent to the reaction zone A in the form of a gas comprising not more than 50% by volume, preferably not more than 40% by volume or not more than 30% by volume, advantageously not more than 25% by volume or not more than 20% by volume, Particularly advantageously, no more than 15% by volume or no more than 10% by volume and more preferably no more than 5% by volume or no more than 2% by volume of other gases (components) (except molecular oxygen). However, since the oxygen demand in the reaction zone A is lower compared to the oxygen demand in the reaction zone B in the process according to the invention, it is used in the process according to the invention, especially in view of economic feasibility. Air is also advantageous as a source of oxygen for the oxygen demand in reaction zone A. From the viewpoint of the minimum gas volume to be transported and compressed, it is preferably supplied to the reaction zone A in the form of pure oxygen. The propane recycle gas (from separation zone IV) recycled to reaction zone A comprises typically only traces of molecular oxygen. The above is applied substantially in the same manner to the supply of molecular oxygen required in the reaction zone B. In other words, the supply of molecular oxygen in the form of pure oxygen is advantageous for the oxygen demand of reaction zone A and reaction zone B, especially since it does not load the gas stream according to the process of the invention with an inert gas which must be recycled again. Uninstallation in another process of the method, except for those that are indispensable.

In the process according to the invention, the typical reaction gas B which can be used to charge the reaction zone B has the following contents: 4 to 25% by volume of propylene, 6 to 70% by volume of propane, 0 to 60% by volume of H ₂ O. , 5 to 60% by volume of O ₂ and 0 to 20, preferably up to 10% by volume of H _2.

Preferred reactant gas B according to the present invention has the following contents: 7 to 25% by volume of propylene, 10 to 40% by volume of propane, 1 to 10% by volume of H ₂ O, 10 to 30% by volume of O ₂ and 0. Up to 10% by volume of H ₂ .

Particularly preferred reactant gas B for use in the above-described packing according to the present invention has the following contents: 15 to 25% by volume of propylene, 20 to 40% by volume of propane, 2 to 8% by volume of H ₂ O, 15 to 30 O _{2 of} volume % and H _{2 of} 0 to 5% by volume.

The advantages of the intermediate propane content in the reaction gas B are, for example, apparent from DE-A 102 45 585.

In the framework of the above composition, when present in a molar ratio of propane in the reaction gas B present in the propylene reaction gas B, V ₁ of the 1 to 9, it is advantageous. In the above composition framework, it is also advantageous when the molar ratio V ₂ of the molecular oxygen present in the reaction gas B to the propylene present in the reaction gas B is from 1 to 2.5. In the context of the present invention, in the above composition framework, when the ratio V ₃ of the molecular hydrogen present in the reaction gas B to the propylene present in the reaction gas B is 0 to 0.80 or 0 to 0.60 or 0.1 to 0.50, It is beneficial. Often, the ratio V ₃ is also from 0.4 to 0.6. In the above composition framework, it is also advantageous when the molar ratio V ₄ of the steam present in the reaction gas B and the total mole amount of propane and propylene present in the reaction gas B is 0 or 0.001 to 10.

Particularly advantageously, the reaction zone B is charged to a reaction gas B, (B, reaction gas starting mixture) of V ₁ to 7 or 1 to 4, or 2 to 6, and especially advantageously 2 to 5 or 3.5 To 4.5. For the reaction gas B starting mixture, it is also preferred when V ₂ is from 1.2 to 2.0 or 1.4 or 1.8. Correspondingly, V ₄ in the starting mixture of reactive gas B is preferably from 0.015 to 5, more preferably from 0.01 to 3, advantageously from 0.01 to 1 and especially advantageously from 0.01 to 0.3 or to 0.1. According to the invention, a non-explosive reactive gas B starting mixture is preferred.

The key to answering the question of whether the starting mixture of reactive gas B is explosive is whether the combustion (ignition, explosion) induced by a local ignition source (such as a white hot platinum wire) is via certain conditions (pressure, temperature). Spreading in the presence of a mixture (refer to the experimental description in DIN 51649 and WO 04/007405). When spread, the mixture should be classified as explosive here. When not spreading, the mixture is classified as non-explosive in this context. When the partially oxidized initial reaction gas mixture of the present invention is non-explosive, it is also generally suitable for use in a reaction gas mixture formed therefrom during partial oxidation (see WO 04/007405).

Quite often, for the process according to the invention, the total content of the components other than propylene, molecular hydrogen, steam, propane, molecular nitrogen and molecular oxygen in the starting mixture of the reaction gas B is usually 20% by volume, or 15% by volume, or 10% by volume, or 5 vol%. Up to 80% by volume of the other components of the starting mixture of the reaction gas B may be ethane and/or methane. In addition, the inclusions may especially be carbon oxides (CO ₂ , CO) and rare gases, but also treat a minor minor component with oxygen, such as formaldehyde, benzaldehyde, methacrolein, acetic acid, propionic acid, Methacrylic acid, etc. Of course, ethylene, isobutylene, n-butane and n-butene are also included in these possible other components of the starting mixture of reactive gas B.

The content of molecular nitrogen in the starting mixture of reactive gas B can vary over a wide range. In principle, the nitrogen content can be up to 70% by volume. In general, the content of molecular nitrogen in the starting mixture of reaction gas B will be 65 vol%, or 60% by volume, or 50% by volume, or 40% by volume, or 30% by volume, or 20% by volume, or 10% by volume, or 5 vol%. In principle, the N ₂ content of the starting mixture of the reaction gas B can be as small as zero. _Minimize total content of B in the reaction gas starting mixture of N so that the gas volume to be transmitted and compressed according to the method of the present invention are minimized.

Conversely, diluting the reaction gas B with molecular nitrogen reduces the formation of propionic acid by-products in the reaction zone B.

In principle, all the processes known in the prior art for the partial dehydrogenation of propane to propylene to propylene are suitable for the heterogeneously catalyzed partial dehydrogenation of propane in reaction zone A, as for example from the literature WO 03/ 076370, WO 01/96271, EP-A 1 17 146, WO 03/011804, EP-A 7 310 77, US-A 3,162, 60, WO 01/96270, DE-A 331 35 73, DE-A 102 45 585, DE-A 103 16 039, DE-A 10 2005 009891, DE-A 10 2005 013039, DE-A 10 2005 022 798, DE-A 10 2005 009 885, DE-A 10 2005 010 111, DE-A 10 2005 049 699, DE-A 10 2004 032 129, DE 10 2005 056377.5, DE 10 2006 017623.5, DE 10 2006 015235.2, DE 10 2005 061626.7 and DE 10 2005 057197.2 and the prior art recognized in these documents ( It is attributable to the fact that the oxidation (combustion) of molecular hydrogen and molecular oxygen in reaction zone A can in principle be followed by any method for the heterogeneously catalyzed partial dehydrogenation of propane. In principle, the heterogeneous catalytic dehydrogenation of propane and the combustion of molecular hydrogen in reaction zone A can also be carried out in different reactors, for example reactors which are connected in series in space. In this case, advantageously, the starting mixture of the reaction gas B in the process according to the invention does not have to comprise molecular oxygen. Quite often, in reaction zone A, the process according to the invention can be carried out only in a single reactor or in, for example, more than one reactor connected in series.

Based on a single pass of propane to reaction zone A through reaction zone A, reaction zone A can in principle be isothermally controlled by controlled heat exchange of a fluid (i.e., liquid or gas) heat carrier directed outside reaction zone A. Configuration. However, based on the same reference, it can also be performed adiabatically, i.e. substantially no controlled heat exchange of the heat carrier used outside of reaction zone A is required. In the latter case, based on the single pass of propane sent to reaction zone A through reaction zone A, the total thermal signature can be endothermic (negative) or self-heating by employing the measures recommended in the above references and described below. Is 0) or exothermic (positive) configuration. It is likewise possible to use all of the catalysts recommended in the above mentioned documents (including prior art recognized in these documents) in the process according to the invention. In principle, the heterogeneous phase in reaction zone A catalyzes the dehydrogenation of propane, regardless of its adiabatic operation and/or isothermal operation (may also change from adiabatic to isothermal in reaction zone A and vice versa), which can be fixed In a bed reactor (or a plurality of (for example) fixed-bed reactors connected in series) or in a moving bed or fluidized bed reactor (the latter is particularly suitable for use by molecular oxygen as it is supplied to the reaction due to its backmixing) The hydrogen in the reaction gas A is burned when the gas A is initially mixed, and the starting mixture of the reaction gas A is heated to the reaction temperature in the reaction zone A).

In general, propane heterogeneously catalyzed partial dehydrogenation to propylene requires a relatively high reaction temperature. The conversion rate of propane can often reach the limit of thermodynamic equilibrium. Typical reaction temperatures are from 300 to 800 ° C or from 400 to 700 ° C, or from 450 to 650 ° C. Dehydrogenation of propane to propylene per molecule produces one molecule of hydrogen. According to the invention, the working pressure in the total reaction zone A is suitably > 1 to 5 bar, preferably 1.5 to 4 bar and advantageously 2 to 3 bar. In principle, the working pressure in reaction zone A can also be a higher or lower value than the values mentioned above. The high temperature and removal of the H ₂ reaction product shifts the equilibrium position of dehydrogenation in reaction zone A toward the propylene to be formed in reaction zone A.

Since the heterogeneous catalytic dehydrogenation reaction proceeds in an increasing volume, the conversion can be increased by lowering the partial pressure of the dehydrogenation product. This can be achieved in a simple manner, for example by dehydrogenation under reduced pressure (however, execution at elevated pressure is generally advantageous for catalyst life) and/or by addition of, for example, steam which typically constitutes an inert gas for the dehydrogenation reaction. This is achieved by essentially inert diluent gas. As a further advantage, the dilution with steam generally causes a reduction in the coking of the catalyst used, since the steam reacts with the formed carbon by the principle of coal gasification. The heat capacity of the steam can also compensate for some of the endothermic properties of dehydrogenation.

Although it has been found that a limited amount of steam has a promoting effect on the activity of a portion of the oxidation catalyst in the downstream reaction zone B, the amount exceeding this is disadvantageous in the reaction zone B for the reasons already mentioned. According to the invention, at least a portion of the hydrogen must additionally be oxidized to steam in reaction zone A. Therefore, according to the present invention, the amount of steam sent to the reaction zone A is measured by the reaction gas A of the catalyst filler sent to the reaction zone A. 20% by volume, preferably 15% by volume and better It is advantageous when it is 10% by volume. However, in general, the amount of steam sent to reaction zone A will generally be 1% by volume, in many cases 2% by volume, or 3 vol% and often 5 vol%.

Other diluents suitable for heterogeneously catalyzing the dehydrogenation of propane are, for example, nitrogen, such as the noble gases of He, Ne and Ar, but also compounds such as CO, CO ₂ , methane and ethane. All of the diluents mentioned may be used alone or in different mixtures. When the above diluent gas is formed as a by-product in the process gas according to the invention, or as a fresh gas or fresh gas component, the process according to the invention requires a discharge in a corresponding manner, which is according to the invention The corresponding fresh gas feed is the reason for the less preference. These possible discharges are present in different separation zones, especially separation zones III and IV.

In principle, suitable dehydrogenation catalysts for heterogeneously catalyzing the dehydrogenation of propane are all of the catalysts known in the prior art. It can be roughly divided into two groups, specifically divided into those having oxidation characteristics (such as chromium oxide and/or aluminum oxide) and divided into deposits on a general oxidation carrier (for example, zirconium dioxide, magnesium oxide, aluminum oxide, dioxide). At least one of ruthenium, titanium dioxide, ruthenium oxide and/or ruthenium oxide is generally the composition of a relatively rare metal such as platinum. Dehydrogenation catalysts which can be used as a result include WO 01/96270, EP-A 731 077, DE-A 10211 275, DE-A 10 131 297, WO 99/46039, US-A 4 788 371, EP-A 705 136, WO 99/29420 US-A 4 220 091, US-A 5 430 220, US-A 5 877 369, EP-A 117 14, 6DE-A 199 37 196, DE-A 199 37 105, US-A 3 670 044, US All of the catalysts recommended in -A 6 566 573, WO 94/29021 and DE-A 199 37 107. In particular, the catalysts of Example 1, Example 2, Example 3 and Example 4 according to DE-A 199 37 107 can be used. Finally, the catalyst of German Application No. 10 2005 044916 should also be recommended as a catalyst for dehydrogenation in reaction zone A. In the process according to the invention, the catalysts can be used as a single catalyst in reaction zone A, since they also generally catalyze the combustion of molecular hydrogen and molecular oxygen.

In principle, the catalytically active bed in reaction zone A can consist solely of catalytically active shaped bodies. However, the catalytically active bed in reaction zone A may of course also consist of a catalytically active shaped body which is diluted with an inert shaped body. Such inert shaped bodies can be produced, for example, from refractory clay (aluminum silicate) or steatite (such as C 220 from CeramTec), or other (preferably substantially void-free) high temperature ceramic materials such as aluminum. Oxide, cerium oxide, titanium dioxide, magnesium oxide, cerium oxide, cerium oxide, mixed zinc or aluminum oxide, cerium oxide, zirconium dioxide, cerium carbide, or other ceric acid salts such as magnesium silicate, and the like a mixture.

The above materials are also suitable for use in the inert top bed and, where appropriate, the end bed of a fixed catalyst bed in accordance with the process of the present invention.

The dehydrogenation catalysts are those catalysts which comprise 10 to 99.9% by weight of zirconium dioxide, 0 to 60% by weight of alumina, ceria and/or titania, and 0.1 to 10% by weight of the first of the periodic table. Or at least one element of the second main group, one element of the third transition group, one element of the eighth transition group, bismuth and/or tin, the limitation being that the sum of the weight percentages is 100% by weight.

Dehydrogenation catalysts useful in the working examples of this document are also particularly suitable.

In a preferred embodiment, the dehydrogenation catalyst includes at least one element of the eighth transition group, at least one element of the first and second main groups, and at least one element of the third and/or fourth main group, and At least one element comprising a third transition group of a lanthanide and a lanthanide element. As an element of the eighth transition group, the active composition of the dehydrogenation catalyst preferably comprises platinum and/or palladium, more preferably platinum. As an element of the first and second main groups, the active composition of the above dehydrogenation catalyst preferably comprises potassium and/or cesium. As an element comprising a third transition group of a lanthanoid element and a lanthanoid element, the active composition of the above dehydrogenation catalyst preferably comprises ruthenium and/or osmium. As an element of the third and/or fourth main group, the active composition of the above dehydrogenation catalyst preferably comprises one or more elements derived from the group consisting of boron, gallium, germanium, antimony, tin and lead, more preferably tin.

Most preferably, the active composition of the above described dehydrogenation catalyst comprises at least one representative of the above-described family of elements in each condition.

Typically, the dehydrogenation catalyst can be a catalyst extrudate (typically from 1 to 10 mm in diameter, preferably from 1.5 to 5 mm; typically from 1 to 20 mm in length, preferably from 3 to 10 mm), in the form of tablets (size) Preferably, the same as the extrudate) and/or the catalyst ring (outer diameter and length are usually 2 to 30 mm or 10 mm in each case, and the wall thickness is suitably 1 to 10 mm or to 5 mm or to 3 mm) ). To perform heterogeneous catalytic dehydrogenation in a fluidized bed (or moving bed), a correspondingly finer powdered catalyst will be used. For reaction zone A, it is preferred to use a fixed catalyst bed in accordance with the present invention.

In principle, there are no restrictions on the geometry of the catalyst to be used for the immobilization of the catalyst bed (especially in the case of supported catalysts) or on the geometry of the inert shaped bodies to be used in the process according to the invention. Particularly commonly used geometries are solid cylinders, hollow cylinders (rings), spheres, cones, pyramids and cubes, as well as extrudates, wheel shapes, stars and monoliths.

The longest dimension of the shaped catalyst body and the inert shaped body (the longest straight line connecting the two points on the surface of the shaped body) may be from 0.5 mm to 100 mm, often from 1.5 mm to 80 mm, and in many cases from 3 mm to 50 Mm or to 20 mm.

In general, the dehydrogenation catalysts (especially the catalysts used in the examples herein and the catalysts recommended in DE-A 19 937 107, especially the exemplary catalysts of the DE-A) enable them to Catalytic dehydrogenation of propane and combustion of molecular hydrogen and combustion of propane/propylene. Compared to the dehydrogenation of propane and its combustion, for example, under competing conditions, hydrogen combustion proceeds more rapidly than the catalysts (ie, it causes, in particular, the lowest combustion of molecular hydrogen under certain conditions). activation energy). The combustion of propane/propylene is usually carried out more rapidly than dehydrogenation over these catalysts.

The above association makes it possible to use only one type of catalyst in the reaction zone A in the process according to the invention.

Advantageously, it should be ensured that in the reaction zone A, whenever the reaction gas A comprises molecular oxygen and starts to contact the above-mentioned catalyst, the reaction gas A comprises at least twice the amount of molecular hydrogen relative to the amount of molecular oxygen present therein, Since the risk of combustion of propane and/or propylene is additionally increased and the converted product of propane in the process according to the invention is reduced, the yield of the target product is reduced.

Of course, the catalyst used to combust molecular hydrogen into water in reaction zone A can also be a catalyst that specifically catalyzes the reaction. Such catalysts are described, for example, in the documents US-A 4,788, 371, US-A 4, 886, 928, US-A 5, 430, 209, US-A 5, 530, 171, US-A 5, 527, 979, and US-A 5, 633, 314.

Suitable reactor types and process variants for performing heterogeneously catalyzed propane dehydrogenation are in principle all of those known in the prior art.

Descriptions of such process variants are found, for example, in all prior art references to dehydrogenation catalysts and in the prior art cited at the outset.

A fairly comprehensive description of a suitable dehydrogenation process in accordance with the present invention also exists in Catalytica Studies Division, Oxidative Dehydrogenation and Alternative Dehydrogenation Processes, Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, California 94043-5272 USA.

Due to the relatively high reaction temperatures required for the heterogeneously catalyzed dehydrogenation of propane, at most a small amount of high boiling, high molecular weight organic compounds comprising carbon are typically formed in reaction zone A of the process according to the invention, which is deposited on the catalyst Surface and therefore inactivated. In order to minimize this adverse concomitant phenomenon, the propane-containing reaction gas A which is intended to cross the catalyst surface at a high temperature of the heterogeneous catalytic dehydrogenation may be appropriately diluted with steam. It may be particularly desirable in the process according to the invention, for example to remove propane from the absorption by means of steam in the separation zone IV by means of steam or by means of a gas comprising steam (for example an inert gas) and A simple way of recycling the propane-laden stripping gas (steam-filled steam) as a propane recycle gas to the reaction zone A. Under conditions that otherwise exist and thus extend the life of the catalyst, the steam will partially or completely remove the carbon deposited by the principles of coal gasification.

Another way to remove the deposited carbon compounds is to, at high temperatures, occasionally pass a gas comprising oxygen (suitably free of hydrocarbons) through the dehydrogenation catalyst (and any catalyst used alone for hydrogen combustion) and thus effectively burn Do as much as possible of the deposited carbon. However, certain inhibition of carbon deposit formation, and thus prolonged catalyst life, is also achieved by adding molecular hydrogen to the propane to be dehydrogenated under heterogeneous catalysis, and then directing it over the dehydrogenation catalyst at elevated temperatures. And as possible. Removing any molecular hydrogen still present in the residual gas present in separation zone II by, for example, upstream of the separation membrane of separation zone III (or outside of separation zone III) in a simple manner, as appropriate, and It is recycled to reaction zone A, which is possible in the process according to the invention. Of course, it is also possible to use molecular hydrogen from another source at this time. For example, molecular hydrogen may also be present (as appropriate) recycled to the residual gas recycle gas in reaction zone A.

Of course, there is also the possibility of adding a mixture of steam and molecular hydrogen to the propane to be dehydrogenated under heterogeneous catalysis.

The addition of molecular hydrogen to the heterogeneous catalytic dehydrogenation of propane also reduces the improper formation of propadiene, propyne and acetylene as by-products.

It allows only fresh propane and propane recycle gas to be sent to reaction zone A in the simplest case to form a reaction gas A to be directed through at least one catalyst bed (herein also referred to as reaction zone A, a filling gas mixture or reaction). Gas A starting mixture). The propane recycle gas may have included a certain amount of molecular oxygen required to cause combustion of hydrogen required in accordance with the present invention in reaction zone A (e.g., when propane is removed from the absorbent in separation zone IV by stripping and steaming) The gas is extracted, for example, when added as molecular oxygen in the form of air. However, the starting mixture of the reaction gas A may also include only fresh propane, a propane recycle gas, and a residual gas recycle gas, in which case the latter may also include molecular oxygen.

Advantageously, the propane recycle gas under the above conditions will comprise just enough steam which, together with the steam formed during the combustion of hydrogen in reaction zone A, can develop its advantageous properties for the overall process. Under the above conditions, there is no need to supply other gas streams to the reaction zone A. The desired conversion in the reaction zone A is carried out by passing the reaction gas A through it a single time.

Of course, the reaction gas A to be directed through the at least one fixed catalyst bed can be formed by adding additional steam and/or additional molecular hydrogen in addition to fresh propane, propane recycle gas and, where appropriate, residual gas recycle gas. Develop its beneficial effects as described in this article. The molar ratio of molecular hydrogen to propane in the filling gas mixture of reaction zone A is usually 5, often 3, in many cases for 1 or 0.1.

The molar ratio of steam to propane in the packed gas mixture of reaction zone A can be, for example, 0 to 30. Suitably it will be from 0.05 to 2 and strongly from 0.1 to 0.5.

In addition, it is also possible to add additional molecular oxygen (in pure form and/or in a mixture with an inert gas) and/or additional inert gas to the filling gas mixture of reaction zone A as needed. The desired conversion of reaction zone A can then be carried out again with a single pass of reaction gas A (filling gas of reaction zone A) without the need to supply other gas streams along the reaction path. In the present context, the reaction route in reaction zone A is understood to mean the flow path of propane through reaction zone A as a function of the dehydrogenation conversion of the propane (conversion in heterogeneous catalytic dehydrogenation). The propane is sent to the reaction gas A before it passes through at least one of the catalyst beds of the reaction zone A.

Suitable reactor configurations for the heterogeneously catalyzed propane dehydrogenation with a single charge of the packed gas mixture through reaction zone A and without intermediate gas feed are, for example, fixed bed tubular reactors or tube bundle reactors. In the reactor, the dehydrogenation catalyst is placed as a fixed bed in a reaction tube or a bundle of reaction tubes. When the combustion system of hydrogen required in the reaction zone A according to the invention is such that the total reaction carried out in the reaction zone A is carried out endothermically, the reaction tube will be suitably heated from the outside according to the invention (it is understood that, if necessary, Can also be cooled). It can be achieved, for example, by burning a gas such as a hydrocarbon such as methane in a space surrounding the reaction tube. Advantageously, the direct form of the catalyst tube is applied to heat only 20 to 30% of the fixed bed and the remaining bed length is heated to the desired reaction temperature via radiant heat released during the combustion process. In this way, a substantially isothermal reaction can be achieved. Suitable for the inner diameter of the reaction tube is about 10 to 15 cm. A typical dehydrogenation tube bundle reactor includes from 300 to 1000 or more reaction tubes. The temperature inside the reaction tube is in the range of 300 to 800 ° C, preferably in the range of 400 to 700 ° C. Advantageously, the starting mixture of reactant gas A is sent to a tubular reactor preheated to the reaction temperature. It is possible that the product gas (mixture) A leaves the reaction tube at a temperature lowering by 50 to 100 °C. However, the outlet temperature can also be higher or at the same level. In the above procedure, an oxidative dehydrogenation catalyst based on chromium oxide and/or alumina is suitably used. An undiluted dehydrogenation catalyst is typically used. On an industrial scale, a plurality of such tube bundle reactors can be operated in parallel and their product gas A can be used in the form of a mixture to fill reaction zone B. Where appropriate, it is also possible that only two of the reactors are present in the dehydrogenation operation and the catalyst packing is regenerated in the third reactor.

The single pass of the filling gas through the reaction zone A can also be effected in a moving bed or fluidized bed reactor, as described, for example, in DE-A 102 45 585 and the literature cited in this document.

In principle, the reaction zone A according to the method of the invention may also consist of, for example, 2 segments. The configuration of the reaction zone A is especially feasible when the charge gas of the reaction zone A does not include any molecular oxygen.

In this case, the actual heterogeneous catalytic dehydrogenation can be achieved in the first section and after the supply of molecular oxygen and/or a mixture of molecular oxygen and inert gas, the hydrogen required according to the invention The homogeneous catalytic combustion can be achieved in the second section.

Quite often, the reaction zone A in the process according to the invention will suitably be operated in this manner in order to pass the total amount of propane to the reaction zone A in a single pass through the reaction zone A. 5 mol% to 60 mol%, preferably 10 mol% to 50 mol%, better 15 mol% to 40 mol%, and the best 20 mol% to 35 mol% is converted in the reaction zone A by dehydrogenation. According to the present invention, the limited conversion in reaction zone A is generally sufficient because the remaining amount of unconverted propane acts substantially as a diluent gas in downstream reaction zone B and is further in the process of the present invention. It can be recycled to reaction zone A substantially without loss. An advantage of the procedure with relatively low propane conversion is that with a single pass of reaction gas A through reaction zone A, the heat required for endothermic dehydrogenation is relatively low and a relatively low reaction temperature is sufficient to effect conversion.

Advantageously, according to the present invention, propellant dehydrogenation (e.g., at a relatively low propane conversion) in the reaction zone A adiabatically (similarly) may be suitable. This means that the charge gas mixture of reaction zone A will typically be first heated to a temperature of from 500 to 700 ° C (or from 550 to 650 ° C) (eg, by direct calcination of the surrounding walls). Typically, a single adiabatic passage through the catalyst bed will then be sufficient to achieve the desired dehydrogenation conversion rate and the combustion of hydrogen required in accordance with the present invention, in the process, depending on the quantitative ratio of the endothermic dehydrogenation of hydrogen to the exothermic combustion, the reactive gas It will be heated, cooled or thermally neutral in the overall assessment. Preferred in accordance with the invention is an adiabatic mode of operation in which the reactant gases are cooled by about 30 to 200 ° C in a single pass. If desired, in the second section of reaction zone A, the hydrogen formed during dehydrogenation can be combusted with the molecular oxygen of the intermediate feed under heterogeneous catalysis. This combustion can likewise be carried out adiabatically.

Significantly, especially in adiabatic operation, a single blast furnace reactor in which the reactant gas A flows axially and/or radially is sufficient for use as a fixed bed reactor.

In the simplest case, it is a single closed reaction volume, such as a container having an inner diameter of 0.1 to 10 m, and even possibly 0.5 to 5 m and in which the fixed catalyst bed is placed on a support device such as a grid. The reaction volume which has been packed with a catalyst and which is substantially insulated during the adiabatic operation is flowed axially by a hot reaction gas A comprising propane. The catalyst geometry can be spherical or toroidal or extrudate. Since the reaction volume can be achieved in this condition by means of an extremely inexpensive device, all catalyst geometries with a particularly low pressure drop are preferred. In particular, there is a catalyst geometry that produces a large cavity volume or in the form of, for example, a monolith or honeycomb structure. To achieve radial flow of the reactive gas A comprising propane, the reactor can, for example, be composed of two concentric cylindrical grids disposed in the outer casing and the catalyst bed can be arranged in its annular gap. In the case of adiabatic conditions, the metal casing (if appropriate) will again be insulated.

Suitable catalyst fillers of the invention for the heterogeneously catalyzed dehydrogenation of propane are also especially the catalysts disclosed in DE-A 199 37 107, in particular all of the catalysts disclosed therein and their catalytic catalysis A mixture of geometrically shaped bodies that are inert to hydrogen.

After prolonging the operating time, the above catalyst can be initially (preferably) nitrogen, for example, in a simple manner, at an inlet temperature of 300 to 600 ° C, often 400 to 550 ° C, in the first regeneration stage. And/or steam diluted air is regenerated across the catalyst bed. The loading of the catalyst (bed) to the regeneration gas (for example, air) may be, for example, 50 to 10000 h ^-1 and the oxygen content of the regeneration gas may be 0.1 or 0.5 to 20% by volume.

In a subsequent further regeneration phase, it is possible to use air as the regeneration gas under otherwise identical regeneration conditions. Suitably from application viewpoint, the recommended catalyst rinsed with an inert gas (e.g. N ₂₎ before the regeneration of the catalyst.

Subsequently, under the same conditions, pure molecular hydrogen or molecular hydrogen diluted with an inert gas (preferably steam and/or nitrogen) should be used. 1% by volume) regeneration is usually feasible.

The heterogeneously catalyzed propane dehydrogenation in reaction zone A according to the process of the invention can be carried out in the same catalyst (bed) loading as the variant with high propane conversion (>30 mol%) (related to the total reaction gas) And in all conditions associated with the propane present in it, at a relatively low propane conversion rate ( 30 mol%) operation. The loading amount of the pair of reactive gases A may be, for example, from 100 to 40 000 or to 10 000 h ^-1 , often from 300 to 7000 h ^-1 , that is, in many cases from about 500 to 4000 h ^-1 . The above loading can also be applied to any catalyst bed specifically used for the combustion of hydrogen in reaction zone A.

In a particularly modest manner, the heterogeneous phase in the reaction zone A catalyzes the dehydrogenation of propane (especially in the case of a single transfer, 15 to 35 mol% propane conversion) can be carried out in a tray reactor, It is the reason that the reaction zone A preferably comprises at least one of the reactors.

The reactor includes more than one catalyst bed for catalytic dehydrogenation in a spatial sequence. The number of catalyst beds may range from 1 to 20, suitably from 2 to 8, but also from 3 to 6. The increased propane conversion rate can be gradually and easily achieved with an increasing number of disks. The catalyst beds are preferably arranged in a radial or axial sequence. Suitably, a fixed catalyst bed type is used in the disc reactor from the application point of view.

In the simplest case, the fixed catalyst beds in the blast furnace reactor are axially aligned or arranged in an annular gap of a concentric cylindrical grid. However, it is also possible to arrange the annular gaps in a section above one another and after the gas has passed radially through a section, it is introduced into a section above or below it.

Suitably, the reaction gas A is passed in a tray reactor over its path from the catalyst bed to the next catalyst bed, intermediate heating, for example by passing it over the surface of the heat exchanger heated by the hot gas (eg The ribs are heated either by passing them through a tube heated by hot combustion gases (and possibly intermediate cooling if desired).

When the tray reactor is additionally insulated, for In the case of a catalyst conversion of 30 mol%, in particular when using the catalysts described in DE-A 199 37 107, in particular the catalysts of the exemplary embodiments, it is particularly sufficient to introduce the reaction gas mixture into a preheating of 450 to 550 ° C. The temperature is in the dehydrogenation reactor (preferably 450 to 500 ° C) and is maintained within this temperature range in the tray reactor. This means that the overall dehydrogenation of propane can thus be achieved at very low temperatures, which has been found to be particularly advantageous for the lifetime of a fixed catalyst bed between regenerations. For higher propane conversion, the reaction gas mixture is suitably introduced into a dehydrogenation reactor preheated to a higher temperature (which can be as high as 700 ° C) and maintained in this high temperature range in the tray reactor.

Even more exaggeratedly, the above intermediate heating is carried out in a direct manner (which allows the self-heating method). To this end, a limited amount of molecular oxygen is before it flows through the first catalyst bed (for example as a component of the residual gas recycle gas and/or the propane recycle gas) (in which case the reaction gas A starting mixture should advantageously be The added molecular hydrogen is included and/or added to the reaction gas A between the downstream catalyst beds. It is therefore possible (usually by itself to catalyze the dehydrogenation catalyst) to produce the combustion of the molecular hydrogen required according to the invention, which is present in the reaction gas A and which has been formed during the heterogeneously catalyzed dehydrogenation of propane and / or added to the reaction gas A in a specific selective and controlled manner (from the application point of view, a catalyst bed packed with a catalyst which specifically burns (specifically) catalytically hydrogen can also be suitably inserted into the disc In the reactor (examples of such an effective catalyst include the documents US-A 4 788 371, US-A 4 886 928, US-A 5 430 209, US-A 5 530 171, US-A 5 527 979 and US-A 5 563 314 of such catalysts; for example, in a tray reactor, the catalyst beds may be recommended to replace the bed comprising the dehydrogenation catalyst; the catalysts are also suitable for the hydrogen in the second section of reaction zone A combustion)). Depending on the amount of molecular hydrogen burned, the heat released by the reaction thus allows for an overall heating mode of operation for heterogeneously catalyzed propane dehydrogenation, or an overall self-heating mode of operation (total thermal characteristics are generally zero) or overall Endothermic mode of operation. When the selected residence time of the reaction gas in the catalyst bed is increased, the propane dehydrogenation may therefore be at a reduced or substantially constant temperature, which makes a particularly long life between 2 regenerations possible and preferred according to the invention. of.

In particular, when molecular hydrogen is burned during the heterogeneous catalytic dehydrogenation of propane (stage 1) and after the heterogeneous catalytic dehydrogenation of propane (stage 2), according to the invention, Indirect heat exchange of components of the reaction gas A charging mixture other than the propane cycle gas and the dehydrogenation cycle gas, and the cooling is carried out in the amount of final hydrogen combustion sent to the reaction zone A after the heterogeneous catalytic dehydrogenation is completed. The reaction gas of reaction zone A (a part can be recycled to the reaction zone A as a dehydrogenation cycle gas as one of the feed streams including propane), and after the final combustion of the molecular hydrogen in the reaction zone A, also by Indirect heat exchange, but for heat exchange with the propane recycle gas, product gas A is cooled prior to its further processing. This is especially true when reaction zone A is configured for overall heat generation.

In general, the oxygen feed as described above should be carried out in this manner in accordance with the invention such that the oxygen content of the reaction gas A is from 0.5 to 50, or to 40, or to 30, based on the amount of molecular hydrogen present therein. %, preferably 10 to 25% by volume. As already mentioned, suitable sources of oxygen comprise pure molecular oxygen (preferably according to the invention) or molecular oxygen diluted with an inert gas such as CO, CO ₂ , N ₂ and/or a noble gas, and also especially air. Preferably, the molecular oxygen system according to the invention comprises not more than 30% by volume, preferably not more than 25% by volume, advantageously not more than 20% by volume, more advantageously not more than 15% by volume, more preferably not more than 10% by volume and more preferably More preferably, no more than 5% by volume of other gases (other than molecular oxygen) are fed in the form of a gas. Particularly advantageously, the oxygen feed is achieved in pure form.

Since 1 mol of hydrogen is burned to H ₂ O to obtain energy about 1 mol of propane dehydrogenated to propylene about twice (about 240 kJ / mol) and H ₂ consumption (about 120 kJ / mol), so as described The self-heating configuration of reaction zone A in the adiabatic disk reactor can be effectively combined with the intended advantages of the process of the present invention, assuming that it requires exactly about 50 moles of molecular hydrogen formed by dehydrogenation in reaction zone A. The burning of the amount of hydrogen.

However, the advantage of the process according to the invention is effective not only when the amount of hydrogen of about 50 mol% of the amount of molecular hydrogen formed in reaction zone A is combusted in reaction zone A. In fact, this advantage is even 25 to 95 mol%, or 100 mol%, or 30 to 90 mol%, or 35 to 85 mol%, or 40 to 80, of the amount of molecular hydrogen formed in the reaction zone A. Mool%, or 45 to 75 mol%, or 50 to 70 mol%, or 35 to 65 mol%, or 40 to 60 mol%, or 45 to 55 mol% of hydrogen is burned in the reaction zone A to obtain water The time (preferably in the above operating mode of the insulated disk reactor) is effective.

In general, the oxygen feed as described above should be carried out so that the oxygen content of the reaction gas A is 0.01 or 0.5 to 3% by volume based on the amount of propane and propylene present therein.

The isothermal catalyzed isothermal dehydrogenation of propane can be further improved by the presence of closed (e.g. tubular) internal components which are advantageously, but not necessarily, filled in the catalyst bed in the tray reactor. The space between the rooms has been exhausted. The internal components can also be placed in a specific catalyst bed. The internals include suitable solids or liquids that evaporate or melt above a certain temperature and that dissipate heat in their manner, and when the temperature falls below this value, condense again and release heat in its execution.

Another way of heating the packed gas mixture for the heterogeneously catalyzed propane dehydrogenation in reaction zone A to the desired reaction temperature is by means of the molecules present in the filling gas mixture as it enters reaction zone A Oxygen (for example, by passing it over a suitable combustion catalyst, for example by simply passing it through and/or over it) and by means of the heat of combustion thus released, burning a portion of the propane and/or H ₂ present therein, heating The reaction temperature to dehydrogenation (this procedure (as already mentioned) is especially advantageous in fluidized bed reactors).

Preferably, in accordance with the invention, reaction zone A will be thermally insulated (externally insulated).

According to the above, the reaction zone A according to the method of the invention can be configured as described in the documents DE-A 10 2004 032 129 and DE-A 10 2005 013 39, but with the difference that the filling gas of the reaction zone A used is The mixture is a mixture of fresh propane, a propane recycle gas comprising steam and, if appropriate, molecular oxygen, a residual gas recycle gas. Reaction zone A is carried out as a (preferably adiabatic) disc reactor wherein the catalyst beds, preferably fixed beds, are arranged in a radial or axial sequence. Advantageously, the number of catalyst beds in the tray reactor is three. Preferably, the heterogeneously catalyzed partial propane dehydrogenation is carried out autonomously. To this end, a limited amount of molecular oxygen or a mixture comprising it and an inert gas is added to the filling gas mixture of reaction zone A, which exceeds the first (fixed) catalyst bed, passes in the flow direction and is in the first Between the (fixed) catalyst bed of the (fixed) catalyst bed (for example, as described in DE 10 2006 017623.5). Therefore, the limited combustion of hydrogen formed during the process of heterogeneously catalyzed propane dehydrogenation (and, if appropriate, the limited combustion of propane and/or propylene, as appropriate) is usually caused by the self-catalysis of the dehydrogenation catalyst. The febrile thermal profile essentially maintains the dehydrogenation temperature.

Suitably, part of the heterogeneously catalyzed dehydrogenation of propane is substantially distributed over three catalyst trays so that the conversion of propane introduced into the reactor is about 20 mol% on a single reactor transfer rate (should be understood In the method according to the invention, it may also be 30 mol%, or 40 mol% or 50 mol% or more. The selectivity for propylene formation is often 90% by mole. The maximum contribution of a single disk to conversion rate shifts from front to back in the flow direction as operating time increases. In general, the catalyst packing is regenerated before the third disk in the flow direction provides the greatest contribution to conversion. Advantageously, the regeneration is achieved when the carbonization of all the disks has reached the same level.

Finally, in the (fixed) catalyst bed arranged downstream of the dehydrogenation, substantial individual combustion of the molecular hydrogen can also be achieved after the intermediate oxygen feed. In principle, it can reach the extent that the product gas A which subsequently leaves the reaction zone A essentially no longer contains any molecular hydrogen.

The total amount of propane and propylene supported on the total amount of catalyst (sum of all beds) 500 l(STP)/l. h and 20 000 l(STP)/l. h (typically 1500 l (STP) / l.h to 2500 l (STP) / l.h), quite favorably the above-mentioned heterogeneously catalyzed partial dehydrogenation of propane. The maximum reaction temperature within the individual fixed catalyst beds is advantageously maintained between 500 ° C and 600 ° C (or to 650 ° C).

A disadvantage of the above method of charging a mixture with a reactive gas A comprising molecular oxygen is that catalytically dehydrogenating all of the propane also catalyzes the combustion of propane and propylene with molecular oxygen (complete oxidation of propane and propylene to carbon oxides and steam) ).

In addition to the measures for adding a gas comprising molecular hydrogen to the reaction gas A filling gas mixture, it can be used according to DE-A 102 11 275 by means of a self-dehydrogenation zone (in reaction zone A) The other part of the molecular hydrogen combustion can be followed by the extraction of the product gas into two parts having the same composition, so that only one of the two parts (if appropriate, part or complete of the molecular hydrogen present therein) After combustion, it is sent to the partial oxidation as product gas A or product gas A ^* , and the other portion is recycled as a component of the reaction gas A (usually a component of the reaction gas A charging gas mixture) to dehydrogenation for equilibration. It is then desirable that the molecular hydrogen present from the dehydrogenation recycle gas from the dehydrogenation itself protects the propane and, where appropriate, the propylene present in the charge gas mixture of reaction zone A from the action of molecular oxygen also present therein. This protection is based on the fact that, as already stated, the combustion of molecular hydrogen, which is generally catalyzed by the same catalyst heterogeneously, into water is preferably kinetic during complete combustion of propane and/or propylene.

According to the teachings of DE-A 102 11 275, the dehydrogenation cycle gas cycle will suitably be achieved by the jet pump principle (which is also known as the circulation mode; the propane cycle gas can act as an actuating jet). The above document also states the possibility of adding molecular hydrogen as another oxidation protection to the filling gas mixture of reaction zone A (which may, for example, be self-derived from separation zone II by means of a separation membrane and still comprise molecular hydrogen The residual gas removes the molecular hydrogen).

The typical reaction gas A charged gas mixture (reaction gas A formed in the reaction zone A) in the process according to the invention has the following contents: 55 to 80% by volume of propane, 0.1 to 20% by volume of propylene, 0 Up to 10% by volume of H ₂ , 0 to 5% by volume of O ₂ , 0 to 20% by volume of N ₂ , and 5 to 15% by volume of H ₂ O.

Preferably, the reactant gas A charging gas mixture (the gas mixture entering the first fixed catalyst bed in the reaction zone A in the flow direction of the reaction gas A) according to the method of the present invention has the following contents: 55 to 80 5% by volume of propane, 0.1 to 20% by volume of propylene, 2 to 10% by volume of H ₂ , 1 to 5% by volume of O ₂ , 0 to 20% by volume of N ₂ , and 5 to 15% by volume of H ₂ O .

The proportion of the components of the reaction gas A filling gas mixture other than the listed components is usually 10% by volume and preferably 6 vol%.

Typical product gas A has the following contents: 30 to 50% by volume of propane, 15 to 30% by volume of propylene, 0 to 10% by volume of H ₂ , preferably 0 to 6% by volume of H ₂ , 10 to 25 volumes % H ₂ O, 0 to 1% by volume of O ₂ , and 0 to 35% by volume of N ₂ .

The proportion of the component of product gas A other than the listed components is usually 10% by volume.

The temperature of the product gas A withdrawn from the reaction zone A is usually from 300 to 800 ° C, preferably from 400 to 700 ° C and more preferably from 450 to 650 ° C.

The pressure of the product gas A leaving the reaction zone A is generally from >1 to 5 bar, preferably from 1.5 to 4 bar and advantageously from 2 to 3 bar.

According to the invention, there is no need to further remove the secondary component (but preferably after mechanical removal of the solid particles present therein according to the teachings of DE-A 10 316 039) or to remove some or all of the condensation (by means of After the steam present in the product gas A is obtained by direct and/or indirect cooling, the product gas A (i.e., as the product gas A ^* ) is used to charge at least one oxidation reactor having the reaction gas B in the reaction zone B.

Regardless of whether condensing removal of water from product gas A is carried out, it is appropriate according to the invention, at least by recycling the gas with propane (and, where appropriate, the other components of the gas mixture of the reaction gas A) (typically excluding the dehydrogenation cycle) The gas)) is thermally exchanged, the hot product gas A is cooled, and then it is further used in the reaction zone B, and thus the propane recycle gas (and, if appropriate, the reaction gas A is charged with other components of the gas mixture).

Further, for at least one oxidation reactor in the reaction zone B charged with the product gas A or the product gas A ^* , the molecular oxygen required to achieve the target in the reaction zone B is added to the product gas A or the product gas. A ^* is sufficient.

As already mentioned, this addition can in principle be pure oxygen or a mixture of molecular oxygen and one or more gases which are chemically inert in reaction zone B (for example N ₂ , H ₂ O, noble gases, CO ₂ ) (for example Air) (also referred to herein as the first source of oxygen) is implemented. Preferably, it comprises, according to the invention, no more than 30% by volume, preferably no more than 25% by volume, advantageously no more than 20% by volume, particularly advantageously no more than 15% by volume, more preferably no more than 10% by volume, and More preferably, it is not more than 5% by volume or not more than 2% by volume of a gas form other than molecular oxygen. Pure oxygen is used very particularly advantageously here.

Generally, the amount of molecular oxygen added is such that the molar ratio of molecular oxygen present to propylene present in the charge gas (reaction gas B starting mixture) of reaction zone B is 1 and 3. A gas is added to the product or a product gas comprising a gas A ^* prior to the molecular oxygen, the temperature of which has been advantageously adjusted to a value of 250 to 370 deg.] C, particularly advantageously of 270 to 320. deg.] C of. A gas system comprising molecular oxygen is present in the feed of reaction gas A or reaction gas A ^* , which is compressed to a pressure generally above the pressure of reaction gas A or A ^* (up to 1 bar if appropriate) to include molecular oxygen The gas can be throttled to the reaction gas A or A ^* in a simple manner (for example, as described in DE-A 103 53 014, when the source of oxygen used is air, it is usually compressed by using a radial compressor. ). The temperature of the gas comprising molecular oxygen generally results in the desired temperature (usually 240 to 300 ° C) at which the reaction gas B is charged to the mixture. In general, the temperature of the gas including molecular oxygen is from 100 to 200 °C.

A typical reaction gas B charging gas mixture includes: 10 to 40% by volume of propane, 5 to 25% by volume of propylene, 0 to 10% by volume of H ₂ , 10 to 30% by volume of O ₂ , and 1 to 10% by volume. H ₂ O.

Preferably, the reaction gas B is filled with a gas mixture comprising: 15 to 40% by volume of propane, 7 to 20% by volume of propylene, 0 to 6% by volume of H ₂ , 15 to 30% by volume of O ₂ , and 1 to 7% by volume. H ₂ O.

In the process according to the invention, the inlet pressure of the reaction gas B into the at least one oxidation reactor of the reaction zone B will generally suitably be from >1 bar to 4 bar, usually from 1.3 bar to 3 bar, in many cases 1.5 to 2.5 bar.

In a manner known per se, the heterogeneous phase catalytic partial phase oxidation of propylene to acrylic acid with molecular oxygen is carried out in principle in two successive steps along the reaction coordinates, wherein the first step produces acrolein and the second step is acrolein. Produces acrylic acid.

The reaction sequence in two successive steps in time is disclosed in a manner known per se to terminate the process according to the invention in the acrolein stage (the stage of predominant acrolein formation) and to carry out the removal of the target product at this stage or to The method of the invention continues to the possibility of primary acrylic acid formation and subsequent removal of only the target product.

When carrying out the process according to the invention to the formation of predominantly acrylic acid, it is advantageous according to the invention to carry out the process in two steps, that is to say in two oxidation stages arranged in series, in which case, in two oxidation stages In each of these, the fixed catalyst bed to be used and, preferably, other reaction conditions such as the temperature of the fixed catalyst bed, are suitably adjusted in an optimum manner.

Although the polymetallic oxide comprising the elements Mo, Fe, Bi, which is particularly suitable as the active composition of the catalyst for the first oxidation stage (propylene → acrolein), can also catalyze the second oxidation stage to some extent (acrolein → acrylic acid) However, generally for the second oxidation stage, the active composition is preferably at least one catalyst comprising a plurality of metal oxides of the elements Mo and V.

A process for the heterogeneously catalyzed partial oxidation of propylene on a fixed catalyst bed to be carried out in reaction zone B according to the invention, thus being particularly suitable as a single-stage process for the preparation of acrolein (and, if appropriate, acrylic acid) or as two The first reaction stage of the stage of preparing acrylic acid, the catalyst of the fixed catalyst bed has at least one multi-metal oxide comprising the elements Mo, Fe and Bi as the active composition.

The use of the reaction gas B obtained according to the invention enables a single-stage heterogeneously catalyzed partial oxidation of propylene to acrolein and, where appropriate, a two-stage heterogeneously catalyzed partial oxidation of acrylic acid or propylene to acrylic acid, as clearly described in document EP-A 70 07 14 (first reaction stage; as described therein, also in a corresponding countercurrent mode of the salt bath and the initial reaction gas mixture through the tube bundle reactor), EP-A 70 08 93 (second reaction stage; as described therein, And also in the corresponding countercurrent mode), WO 04/085369 (especially the document is considered as an integral part of this document) (as a two-stage process), WO 04/85363, DE-A 103 13 212 (first reaction stage), EP -A 1 159 248 (as a two-stage process), EP-A 1 159 246 (second reaction stage), EP-A 1 159 247 (as a two-stage process), DE-A 199 48 248 (as a two-stage process) , DE-A 101 01 695 (single or two stages), WO 04/085 368 (as a two-stage method), DE-A 10 2004 021 764 (two stages), WO 04/085 362 (first reaction stage) , WO 04/085 370 (second reaction stage), WO 04/085 365 (second reaction stage), WO 04/085367 (two stages), EP-A This is carried out as described in 990 636, EP-A 1 007 007 and EP-A 1 106 598.

It is especially true for all working examples contained in these documents. It can be carried out as described in the literature, but the difference is that the starting reaction gas mixture for the first reaction stage (propylene to acrolein) is the reaction gas B produced according to the invention. With regard to the remaining parameters, the procedure is the same as in the working examples of the documents mentioned (especially with respect to the reactant loading of the fixed catalyst bed and the fixed catalyst bed). When the procedure in the above working examples of the prior art is two-stage and there is a second oxygen feed between the two reaction stages (in the second preferred air feed according to the less preferred mode of the invention), proceed in an appropriate manner The molar ratio of molecular oxygen to acrolein in the filler mixture adjusted to achieve the second reaction stage corresponds to the Perm ratio in the working examples of the documents mentioned.

Advantageously, according to the invention, the amount of oxygen in reaction zone B is such that product gas B still comprises unconverted molecular oxygen (suitably From 0.5 to 6% by volume, advantageously from 1 to 5% by volume, preferably from 2 to 4% by volume. In the case of a two-stage procedure, the above applies to each of the two oxidation stages.

Multimetal oxide catalysts which are particularly suitable for a particular oxidation stage have been previously described many times and are well known to those skilled in the art. For example, EP-A 253 409 on page 5 refers to the corresponding US patent.

The advantageous catalysts for the particular oxidation stage are also disclosed by DE-A 44 31 957, DE-A 10 2004 025 445 and DE-A 44 31 949. It is especially true for those having the general formula I in the two previously mentioned documents. Particularly advantageous catalysts for the particular oxidation stage are known from the documents DE-A 103 25 488, DE-A 103 25 487, DE-A 103 53 954, DE-A 103 44 149, DE-A 103 51 269, DE-A 103 50 812 and DE-A 103 50 822 disclose.

For the reaction stage of the invention in which the heterogeneous phase catalyzed partial phase oxidation of propylene to acrolein or acrylic acid or a mixture thereof, the applicable multimetal oxide composition is in principle all of the active compositions including Mo, Bi and Fe. Metal oxide composition.

The multimetal oxide active compositions of the general formula I of DE-A 199 55 176, the polymetal oxide active compositions of the general formula I of DE-A 199 48 523, DE- A multi-metal oxide active composition of the general formula I, II and III of A 101 01 695, a polymetal oxide active composition of the general formula I, II and III of DE-A 199 48 248, and DE-A 199 55 A multimetal oxide active composition of the formula I, II and III of 168, and a multimetal oxide active composition as specified in EP-A 700 714.

Also suitable for this reaction stage are multimetal oxide catalysts including Mo, Bi and Fe disclosed in the following documents: Research Disclosure No. 497012 of August 29, 2005 (unintentionally, in its working example) , the specific surface area is incorrectly expressed in cm ² /g and is not correctly reported as m ² /g with the same value), DE-A 100 46 957, DE-A 100 63 162, DE-C 3 338 380, DE-A 199 02 562, EP-A 15 565, DE-C 2 380 765, EP-A 8 074 65, EP-A 27 93 74, DE-A 330 00 44, EP-A 575 897, US-A 4438217, DE- A 19855913, WO 98/24746, DE-A 197 46 210 (to those of the general formula II), JP-A 91/294239, EP-A 293 224 and EP-A 700 714. It is especially suitable for exemplary embodiments in such documents, and among these multimetal oxide catalysts, EP-A 15 565, EP-A 575 897, DE-A 197 46 210 and DE-A are particularly preferred. 198 55 913 of them. Particularly preferred herein is a catalyst according to example 1c from EP-A 15 565 and intended to be prepared in a corresponding manner but having an active composition having the composition Mo ₁₂ Ni _6.5 Zn ₂ Fe ₂ Bi ₁ P _0.0065 K _0.06 O _x . 10 SiO ₂ catalyst. The focus is also from DE-A 198 55 913, as an example of a non-loaded hollow cylindrical catalyst with a geometry of 5 mm × 3 mm × 2 mm (outer diameter × height × inner diameter) with sequence number 3 (stoichiometry: Mo ₁₂ Co ₇ Fe ₃ Bi _0.6 K _0.08 Si _1.6 O _x ), and an unsupported multimetal oxide II catalyst according to Example 1 of DE-A 197 46 210. Mention should also be made of the metal oxide catalysts of US-A 4,438,217. When the hollow cylinders have 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 The latter is especially true when the geometry is mm x 4 mm (each OD x height x ID). Other possible catalyst geometries in this context are extrudates (e.g., 7.7 mm in length and 7 mm in diameter; or 6.4 mm in length and 5.7 mm in diameter).

A plurality of multimetal oxide active compositions suitable for the steps from propylene to acrolein and, if appropriate, acrylic acid, may be encompassed by Formula IV: Mo ₁₂ Bi _a Fe _b X ¹ _c X ² _d X ³ _e X ⁴ _f O _n (IV)

The variables are defined as follows: X ¹ = nickel and / or cobalt, X ² = bismuth, an alkali metal and / or an alkaline earth metal, X ³ = zinc, phosphorus, arsenic, boron, antimony, tin, antimony, lead and / Or tungsten; X ⁴ = yttrium, 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 2, f = 0 to 10 and n = the number determined by the valence and frequency of the element other than oxygen in the IV.

It can be obtained in a manner known per se (see, for example, DE-A 4 023 239) and is generally substantially shaped to obtain spheres, rings or cylinders, or as a coated catalyst, ie a coating of the active composition. It is used in the form of a shaped inert carrier. Of course, it can also be used as a catalyst in powder form (for example, in a fluidized bed reactor).

In principle, the active composition of the formula IV can, in a simple manner, obtain a very fine, preferably finely powdered dry mixture having a composition corresponding to its stoichiometry, from a suitable source of its basic components, and at 350 It was prepared by calcining it at a temperature of 650 °C. The calcination can be carried out under an inert gas or under an oxidizing atmosphere such as air (mixture of inert gas and oxygen) and in a reducing atmosphere (for example, a mixture of inert gas, NH ₃ , CO and/or H ₂ ). The calcination time can range from a few minutes to a few hours, and generally decreases with temperature. Suitable sources of the basic components of the multimetal oxide active composition IV are those compounds which have become oxides, and/or which can be converted to oxides by heating at least in the presence of oxygen.

In addition to the oxides, such suitable starting compounds include, in particular, 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 COOH, NH ₄ CH ₃ CO ₂ and/or ammonium oxalate may additionally be incorporated into the finely dried mixture, which The compound is decomposed and/or may be decomposed at the latest calcination at a later time to obtain a compound which is released in a gaseous form).

The starting compound for preparing the multimetal oxide active composition IV can be finely mixed in a dry form or in a wet form. When it is mixed in a dry form, the starting compound is suitably used as a fine powdery powder and calcined after mixing and optionally compression. However, it is preferred to finely mix in a wet form. Typically, the starting compounds are admixed with each other in the form of aqueous solutions and/or suspensions. When the starting material is the sole source of the basic component in dissolved form, especially fine dry mixtures are obtained in the mixing process. The solvent used is preferably water. Subsequently, the resulting aqueous composition is dried, and the drying method is preferably carried out by spray drying the aqueous mixture at an outlet temperature of 100 to 150 °C.

The multimetal oxide active composition of formula IV can be used in the form of a powder or shaped into a specific catalyst geometry for the "propylene to acrolein (and, if appropriate, acrylic)" step, and the shaping can be effected before or after the final calcination. For example, the unsupported catalyst can be derived from the active composition in powder form or its uncalcined and/or partially calcined precursor composition by compression into the desired catalyst geometry (eg, by tableting or extrusion). And, if appropriate, an additive such as graphite or stearic acid as a lubricant and/or a forming aid and a reinforcing agent such as glass microfiber, asbestos, tantalum carbide or potassium titanate are added. It is also possible to use hexagonal boron nitride as an auxiliary in forming to replace graphite, as recommended by DE-A 10 2005 037 678. Examples of suitable non-supported catalyst geometries include solid cylinders or hollow cylinders having an outer diameter and length of 2 to 10 mm. In the case of a hollow cylinder, a wall thickness of 1 to 3 mm is advantageous. Of course, the unsupported catalyst may also have a spherical geometry and may have a spherical diameter of 2 to 10 mm.

Especially in the case of unsupported catalysts, a particularly advantageous hollow cylindrical geometry is 5 mm x 3 mm x 2 mm (outer diameter x length x inner diameter).

Of course, the powdered active composition or its powdered precursor composition which is still to be calcined and/or partially calcined can also be formed by application to a preformed inert catalyst support. The coating of the carrier body to produce a coated catalyst system is typically carried out in a suitable rotatable container, as disclosed, for example, in DE-A 29 09 671, EP-A 293 859 or EP-A 714 700. To coat the carrier body, the powder composition to be coated is suitably wetted and, after coating, dried again, for example by means of hot air. The coating thickness of the powder composition applied to the carrier body is suitably selected in the range of 10 to 1000 μm, preferably in the range of 50 to 500 μm, and more preferably in the range of 150 to 250 μm.

Suitable carrier materials are conventional porous or non-porous aluminum oxides, cerium oxide, cerium oxide, zirconium dioxide, tantalum carbide or ceric acid salts such as magnesium silicate or aluminum silicate. With regard to the target reaction on which the process according to the invention is based, it generally appears to be substantially inert. Although preferably a regular shaped carrier body having a significant surface roughness, such as a sphere or hollow cylinder, the carrier body can have a regular or irregular shape. Suitable carrier bodies are substantially non-porous, surface roughened spherical supports made from block talc having a diameter of from 1 to 10 mm or to 8 mm, preferably from 4 to 5 mm. However, suitable carrier bodies are also cylindrical, having a length of 2 to 10 mm and an outer diameter of 4 to 10 mm. In the case of a ring suitable as a carrier body according to the invention, the wall thickness is also usually from 1 to 4 mm. The annular carrier body to be preferably used according to the invention has a length of 2 to 6 mm, an outer diameter of 4 to 8 mm and a wall thickness of 1 to 2 mm. A carrier body suitable according to the invention is also in particular a ring of geometric shape 7 mm x 3 mm x 4 mm (outer diameter x length x inner diameter). Of course, the fineness of the catalytically active oxide composition to be applied to the surface of the carrier body is adapted to the desired coating thickness (see EP-A 714 700).

The multimetal oxide active composition to be used in the step of propylene to acrolein (and, if appropriate, acrylic acid) is also a composition of the 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)

The variables are each defined as follows: Y ¹ = only yttrium, or at least one of lanthanum and lanthanum, cerium, tin and copper, Y ² = molybdenum, or tungsten, or molybdenum and tungsten, Y ³ = an alkali metal, lanthanum And/or 钐, Y ⁴ = an alkaline earth metal, nickel, cobalt, copper, manganese, zinc, tin, cadmium and/or mercury, Y ⁵ = iron, or at least one of iron and elemental chromium and bismuth, Y ⁶ = phosphorus, arsenic, boron and/or antimony, Y ⁷ = a rare earth metal, titanium, zirconium, hafnium, tantalum, niobium, tantalum, niobium, silver, gold, aluminum, gallium, indium, antimony, bismuth, lead, antimony 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' = the number determined by the atomic valence and frequency of the element other than oxygen in V and p, q = p/q ratio of 0.1 to 10 _a three-dimensional region comprising a chemical composition Y ¹ _a' Y ² _b' O _{x '} , which is delimited from its local environment and its largest diameter (through the center of the region and the connected region due to its composition different from its local environment) The longest line of the two points on the surface (interface) is 1 nm to 100 μm, often 10 nm to 500 nm or 1 μm to 5 0 or 25 μm.

Polymetallic oxide compositions V which are particularly advantageous in accordance with the present invention are those in which Y ¹ is only ruthenium.

Preferred in these compositions are those of the 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)

The variables are each defined as follows: Z ² = molybdenum, or tungsten, or molybdenum and tungsten, Z ³ = nickel and / or cobalt, Z ⁴ = bismuth, 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" = the number determined by the atomic valence and frequency of the oxygen-excluding element in VI, p", q" = its p"/q" ratio is 0.1 to 5, preferably 0.5 to 2, Very particularly preferred are compositions VI wherein Z ² _b" = (tungsten) _b" and Z ² ₁₂ = (molybdenum) ₁₂ .

[Y ^{1] of a} multimetal oxide composition V (polymetal oxide composition VI) suitable according to the invention in a multimetal oxide composition V (polymetal oxide composition VI) suitable according to the invention At least 25 mol% (preferably at least 50 mol% and more preferably at least 100 mol%) of the total proportion of _a' Y ² _b' O _x' ] _p ([Bi _a" Z ² _b" O _x" ] _p" ) Is a chemical composition Y ¹ _a' Y ² _b' O _x' [Bi _a which is delimited from its local environment due to its chemical composition different from its local environment and whose maximum diameter is in the range of 1 nm to 100 μm [Bi _a It is also advantageous when the three-dimensional region form of _" Z ² _b" O _x" ].

With regard to forming, the statement made for the multimetal oxide composition IV catalyst is applicable to the multimetal oxide composition V catalyst.

The preparation of the multi-metal oxide active composition V is described in, for example, EP-A 575 897 and DE-A 198 55 913.

The inert carrier materials recommended above are also (particularly) suitable for use as an inert material for the dilution and/or delimitation of a suitably fixed catalyst bed, or as an upstream bed for protecting and/or heating a gas mixture.

For the second step (second reaction stage) in which the acrolein heterogeneously catalyzes the partial oxidation of the gas phase to acrylic acid, as already stated, the suitable active composition of the desired catalyst is in principle all metals containing Mo and V. Oxide compositions, such as those in DE-A 100 46 928.

Most of them (for example, compositions of DE-A 198 15 281) can be covered by the 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 )

The variables are each defined as follows: 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 = the number determined by the valence and frequency of the element other than oxygen in VII.

Active multimetal oxides VII in accordance with the preferred embodiment of the present invention as defined by the following general formula VII the variables of their contemplated embodiments: 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 = determined by the atomic valence and frequency of the element other than oxygen in VII number.

However, very particularly preferred multimetal oxide VII according to the invention is an oxide of the formula VIII: Mo ₁₂ V _a' Y ¹ _b' Y ² _c' Y ⁵ _f' Y ⁶ _g' O _{n '} (VIII )

Wherein 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' = the number determined by the valence and frequency of the element other than oxygen in VIII.

Suitable multimetal oxide active compositions (VII) according to the invention are obtainable in a manner known per se (for example as disclosed in DE-A 43 35 973 or in EP-A 714 700).

In principle, the multimetal oxide active compositions suitable for the "acrolein to acrylic acid" step, in particular the compositions of the formula VII, can be obtained in a simple manner by corresponding sources from their basic components. The stoichiometric composition is prepared in a very fine, preferably finely powdered dry mixture and calcined at a temperature of from 350 to 650 °C. The calcination can be carried out under an inert gas or under an oxidizing atmosphere such as air (mixture of inert gas and oxygen) and in a reducing atmosphere (for example, inert gas and reduction such as H ₂ , NH ₃ , CO, methane and/or acrolein). The gas is carried out under a mixture of a gaseous gas or a reducing gas mentioned by itself. The calcination time can range from a few minutes to a few hours, and generally decreases with temperature. Suitable bases for the basic components of the multimetal oxide active composition VII are those compounds which have become oxides and/or which can be converted to oxides by heating at least in the presence of oxygen.

The starting compound for preparing the multimetal oxide active composition VII can be finely mixed in a dry form or in a wet form. When it is mixed in a dry form, the starting compound is suitably used in the form of a finely powdered powder and calcined after mixing and, where appropriate, compression. However, it is preferred to finely mix in a wet form.

It is usually carried out by mixing the starting compounds in the form of aqueous solutions and/or suspensions with one another. When the starting material is the sole source of the basic component in dissolved form, especially fine dry mixtures are obtained in the mixing process. The solvent used is preferably water. Subsequently, the resulting aqueous composition is dried, and the drying method is preferably carried out by spray drying the aqueous mixture at an outlet temperature of 100 to 150 °C.

The resulting multimetal oxide compositions, especially those of the formula VII, may be used in acrolein oxidation in powder form (for example in a fluidized bed reactor) or shaped into a specific catalyst geometry, and the forming may be It is carried out before or after the final calcination. For example, the unsupported catalyst can be added from the active composition in powder form or its uncalcined precursor composition by compression to the desired catalyst geometry (eg, by tableting or extrusion), as appropriate, for example. It is prepared by using a lubricant or a stearic acid auxiliaries and/or a forming aid and a reinforcing agent such as glass microfiber, asbestos, tantalum carbide or potassium titanate. An example of a suitable unsupported catalyst geometry is a solid cylinder or hollow cylinder having an outer diameter and length of 2 to 10 mm. In the case of a hollow cylinder, a wall thickness of 1 to 3 mm is advantageous. Of course, the unsupported catalyst may also have a spherical geometry, in which case the spherical diameter may be 2 to 10 mm (e.g., 8.2 mm or 5.1 mm).

Of course, the powdery active composition or powdery precursor composition which is still to be calcined can also be formed by application to a preformed inert catalyst support. The coating of the carrier body to prepare the coated catalyst is typically carried out in a suitable rotatable container, as disclosed, for example, in DE-A 2 909 671, EP-A 293 859 or EP-A 714 700.

To coat the carrier body, the powder composition to be coated is suitably wetted and, after coating, dried again, for example by means of hot air. The coating thickness of the powder composition applied to the carrier body is suitably selected in the range of 10 to 1000 μm, preferably in the range of 50 to 500 μm, and more preferably in the range of 150 to 250 μm.

Suitable carrier materials are conventional porous or non-porous aluminum oxides, cerium oxide, cerium oxide, zirconium dioxide, tantalum carbide or ceric acid salts such as magnesium silicate or aluminum silicate. Although preferred is a regularly shaped carrier body having a significant surface roughness, such as a sphere or hollow cylinder having a coarse sand layer, the carrier body can have a regular or irregular shape. Suitable carrier bodies comprise substantially non-porous, surface roughened spherical supports made from block talc having a diameter of from 1 to 10 mm or to 8 mm, preferably from 4 to 5 mm. In other words, a suitable spherical geometry can have a diameter of 8.2 mm or 5.1 mm. However, suitable carrier bodies also include cylinders having a length of 2 to 10 mm and an outer diameter of 4 to 10 mm. In the case of a ring as the main body of the carrier, the wall thickness is also usually from 1 to 4 mm. The annular carrier body to be used preferably has a length of 2 to 6 mm, an outer diameter of 4 to 8 mm and a wall thickness of 1 to 2 mm. Suitable carrier bodies are also in particular rings of geometric shape 7 mm x 3 mm x 4 mm (outer diameter x length x inner diameter). Of course, the fineness of the catalytically active oxide composition to be applied to the surface of the carrier body is suitable for the desired coating thickness (refer to EP-A 714 700).

The advantageous multi-metal oxide active composition to be used in the "acrolein to acrylic acid" step is also a composition of the formula IX [D] _p [E] _q (IX)

The variables are each defined as follows: 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 x", y" = determined by the atomic valence and frequency of the element other than oxygen in IX And p, q = the p/q ratio is from 160:1 to 1:1 divided by zero, and the composition can be obtained by the following method: separately combining the multi-metal oxide in a fine powder form E (Starting Composition 1) preformed Z ⁷ ₁₂ Cu _h" H _i" O _y" (E)

And then incorporating the preformed solid starting composition 1 into the aqueous solution, the aqueous suspension or incorporating the elements Mo, V, Z ¹ , Z ² , Z ³ , Z ⁴ , Z ⁵ , at the desired p:q ratio, In a finely powdered dry mixture of the source of Z ⁶ (starting composition 2), the mixture comprises the above elements in stoichiometric D: Mo ₁₂ V _a" Z ¹ _b "Z ² _c" Z ³ _d" Z ⁴ _{e "} Z ⁵ _f" Z ⁶ _g " (D)

The resulting aqueous mixture is dried and the resulting dried precursor composition is calcined at a temperature of from 250 to 600 ° C before or after drying to obtain the desired catalyst geometry.

Preferred is a multimetal oxide composition IX wherein the preformed solid starting composition 1 is incorporated into the aqueous starting composition 2 at a temperature of <70 °C. A detailed description of the preparation of the multimetal oxide composition VI catalysts is described, for example, in EP-A 668 104, DE-A 197 36 105, DE-A 100 46 928, DE-A 197 40 493 and DE-A 195 28 646. in. With regard to forming, the statements made for the multimetal oxide composition VII catalyst are applicable to the multimetal oxide composition IX catalyst.

The multimetal oxide catalysts which are suitable for the "acrolein to acrylic acid" step are also the catalysts of DE-A 198 15 281, in particular the catalysts having the polymetal oxide active compositions of the formula I herein.

Advantageously, the unsupported catalyst ring is used in the step from propylene to acrolein and the coated catalyst ring system is used in the step from acrolein to acrylic acid.

The partial oxidation of propylene to acrolein (and, if appropriate, acrylic acid) according to the process of the invention can be carried out, for example, as described in DE-A 44 31 957, for example in a single-zone multi-catalyst tube fixed bed reactor, The catalyst is carried out. In this case, the reaction gas mixture and the heat carrier (heat exchange medium) can be observed in the cocurrent or countercurrent flow through the reactor.

The reaction pressure is usually in the range of 1 to 3 bar and the total space velocity of the fixed catalyst bed of the reaction gas B is preferably 1500 to 4000 or 6000 l (STP) / l. h or more. The propylene loading (the propylene hourly space velocity on the fixed catalyst bed) is usually 90 to 200 l (STP) / l. h or to 300 l(STP)/l. h or more. According to the invention, greater than 135 l (STP) / l. h or 140 l(STP)/l. h or 150 l(STP)/l. h or 160 l(STP)/l. The propylene loading of h is particularly preferred because the initial reaction gas mixture of the invention of reaction zone B causes advantageous hot spot behavior due to the presence of unconverted propane and, where appropriate, molecular hydrogen (all of the above applications do not take into account the fixed bed) Specific choice of reactor).

The charge gas mixture preferably flows from above to a single zone multiple catalyst tube fixed bed reactor. The heat exchange medium used is suitably a salt melt, preferably 60% by weight of potassium nitrate (KNO ₃ ) and 40% by weight of sodium nitrite (NaNO ₂ ), or 53% by weight of potassium nitrate (KNO ₃ ), 40% by weight of sodium nitrite (NaNO ₂ ) and 7% by weight of sodium nitrate (NaNO ₃ ).

Viewed through the reactor, as already described, the salt melt and the reaction gas mixture can be co-current or directed in countercurrent. The salt melt itself is preferably guided around the catalyst tube in a curved manner.

When flowing from top to bottom to the catalyst tube, the catalyst tube is suitably packed from bottom to top with the following catalyst (for bottom-up flow, the loading sequence is suitably reversed): - First, 40 to 80 of the length of the catalyst tube % or up to 60% of the length, only the catalyst or a mixture of catalyst and inert material, the latter constituting up to 30 or up to 20% by weight by weight of the mixture (C section); - thereafter, 20 of the total length of the manifold To a length of 50% or to 40%, only a catalyst or a mixture of a catalyst and an inert material, the latter constituting a weight ratio of up to 40% by weight (B section) in terms of a mixture; and - finally, 10 to the length of the manifold For a length of 20%, a bed of inert material (section A) is preferably selected so that it produces a very small pressure drop.

The C segment is preferably undiluted.

When the catalyst used is a catalyst according to the research disclosure No. 497012 of August 29, 2005 or according to Example 1 of DE-A 100 46 957 or Example 3 according to DE-A 100 46 957, and the inert substance used is The above-described loading variant is particularly suitable when having a block talc having a geometric shape of 7 mm x 7 mm x 4 mm (outer diameter x height x inner diameter). Regarding the salt bath temperature, the statement of DE-A 4 431 957 applies.

However, the partial oxidation of propylene to acrolein (and, if appropriate, acrylic acid) in the reaction zone B can also be carried out as described in DE-A 199 10 506, DE-A 10 2005 009 885, DE-A 10 2004 032 129, DE- A 10 2005 013 039 and DE-A 10 2005 009 891 and DE-A 10 2005 010 111 are carried out, for example, in a two-zone multi-catalyst tube fixed bed reactor using the catalyst. In both of the above conditions (and the process according to the invention is quite common), the conversion of propylene achieved by a single transfer is usually 90 mol% or 95 mol% value, and the selectivity of acrolein formation A value of 90 mol%. According to the invention, the partial oxidation of the propylene of the invention to acrolein, or acrylic acid or a mixture thereof is described in EP-A 1 159 244 and is preferably as described in WO 04/085 363 and WO 04/085 362. achieve.

The documents EP-A 1 159 244, WO 04/085 363 and WO 04/085 362 are considered as an integral part of this document.

In other words, partial oxidation of propylene to acrolein (and, if appropriate, acrylic acid) to be carried out in reaction zone B can be carried out particularly advantageously on a fixed catalyst bed having an increased propylene loading and at least two temperature zones.

In this regard, reference is made to, for example, EP-A 1 159 244 and WO 04/085 362.

In the case where the two-stage partial oxidation of propylene to acrolein, the second step is carried out, that is, the partial oxidation of acrolein to acrylic acid, as described in DE-A 44 31 949, for example, in a single-zone multi-catalyst tube In the bed reactor, the catalyst is used. In this reaction stage, the reaction gas mixture and the heat carrier can be observed in the same direction via a reactor. In general, the foregoing partial oxidation of the propylene of the present invention to acrolein is, in principle, as indicated (where appropriate, after intermediate cooling (which may be indirectly or directly by, for example, a second oxygen addition (or In the second air addition according to the less preferred mode of the invention, it is achieved, that is to say that it is not necessary to remove the secondary component, which is introduced into the second reaction stage, that is to say into the partial oxidation of acrolein.

In the second step, molecular oxygen required for partial oxidation of acrolein may already be present in the starting mixture of the reaction gas B in which the propylene is partially oxidized to acrolein. However, it may also be added partially or completely directly to the first reaction stage, ie partial oxidation of propylene to a product gas mixture of acrolein (which may be carried out in the form of a second air, but preferably in the form of pure oxygen or an inert gas with oxygen) (better 50% by volume, or 40% by volume, or 30% by volume, or 20% by volume, or 10% by volume, or 5 vol%, or 2% by volume of the mixture is achieved). For the pressure and temperature of the second oxygen source, the same pressure and temperature as stated for the first oxygen source are applied. Regardless of the procedure, the loading gas mixture (initial reaction gas mixture) in which the acrolein is partially oxidized to acrylic acid advantageously has the following contents: 3 to 25% by volume of acrolein, 5 to 65% by volume of molecular oxygen, 6 to 70% by volume of propane, 0 to 20% by volume of molecular hydrogen, 8 to 65% by volume of steam and 0 to 70% by volume of molecular nitrogen.

The above starting reaction gas mixture preferably has the following contents: 5 to 20% by volume of acrolein, 5 to 15% by volume of molecular oxygen, 10 to 40% by volume of propane, 0 to 10% by volume of molecular hydrogen and 5 Up to 30% by volume of steam.

The above starting reaction gas mixture preferably has the following contents: 10 to 20% by volume of acrolein, 7 to 15% by volume of molecular oxygen, 20 to 40% by volume of propane, 3 to 10% by volume of molecular hydrogen and 15 Up to 30% by volume of steam.

The nitrogen content of the above mixture will usually be 20% by volume, preferably 15% by volume, better 10% by volume, and the best 5 vol%. The molar ratio of O ₂ and acrolein present in the filling gas mixture of the second oxidation stage, O ₂ : acrolein, advantageously according to the invention is generally 0.5 and 2, often for And 1.5.

The CO ₂ content of the filling gas mixture in the second oxidation stage will usually be 5 vol%.

When in the first reaction stage (propylene → acrolein), the reaction pressure in the second reaction stage (acrolein → acrylic acid) is also usually in the range of >1 to 4 bar, preferably in the range of 1.3 to 3 bar. In many cases, in the range of 1.5 to 2.5 bar, and the total space velocity of the (starting) reaction gas mixture on the fixed catalyst bed is preferably 1500 to 4000 or to 6000 l (STP) / l. h or more. The acrolein loading (the acrolein hourly space velocity on the fixed catalyst bed) is usually 90 to 190 l (STP) / l. h or to 290 l(STP)/l. h or more. More than 135 l(STP)/l. h, or 140 l(STP)/l. h, or 150 l(STP)/l. h, or 160 l(STP)/l. The acrolein loading of h is particularly preferred because the initial reaction gas mixture to be used in accordance with the present invention also produces advantageous hot spot behavior due to the presence of propane and, where appropriate, molecular hydrogen.

The temperature of the filling gas mixture in the second oxidation stage is usually from 220 to 270 °C.

The acrolein conversion ratio of the reaction gas mixture in a single pass through the fixed catalyst bed of the second oxidation stage is suitably 90 mol% and the accompanying selectivity of acrylic acid formation 90 mol%.

The charge gas mixture is also preferably flowed from above to a single zone multiple catalyst tube fixed bed reactor. The heat exchange medium used in the second stage is also suitably a salt melt, preferably 60% by weight of potassium nitrate (KNO ₃ ) and 40% by weight of sodium nitrite (NaNO ₂ ), or 53% by weight of nitric acid. Potassium (KNO ₃ ), 40% by weight of sodium nitrite (NaNO ₂ ) and 7% by weight of sodium nitrate (NaNO ₃ ). Viewed through the reactor, as already described, the salt melt and the reaction gas mixture can be co-current or directed in countercurrent. The salt melt itself is preferably guided around the catalyst tube in a curved manner.

When flowing from top to bottom to the catalyst tube, the catalyst tube is suitably packed from bottom to top as follows: first, for a length of 50 to 80 or 70% of the length of the catalyst tube, only a catalyst or a mixture of a catalyst and an inert material, the latter In a proportion by weight of the mixture of up to 30 (or 20)% by weight (C-zone); thereafter, for a length of from 20 to 40% of the length of the manifold, only a catalyst or a mixture of catalyst and inert material, the latter being a mixture Constituting a weight ratio of up to 50 or up to 40% by weight (B section); and finally, for a length of 5 to 20% of the length of the manifold, a bed of inert material (A section), which is preferably selected so that Produces a very small pressure drop.

The C segment is preferably undiluted. For example, the heterogeneous phase of acrolein catalyzes the partial oxidation of the gas phase to acrylic acid (especially in the high acrolein loading on the catalyst bed and the high vapor content of the filling gas mixture), the B segment can also be composed of two consecutive Catalyst dilution composition (in order to minimize hot spot temperature and hot spot temperature sensitivity). From bottom to top, up to 30 (or 20)% by weight of inert material is first used and subsequently >20% to 50 or 40% by weight of inert material. The C segment is then preferably undiluted.

For bottom-up flow to the catalyst tube, the catalyst tube loading is suitably reversed.

When the catalysts used are the catalysts according to the preparation example 5 of DE-A 100 46 928 or the catalysts according to DE-A 198 15 281 and the inert materials used have a geometry of 7 mm×7 mm×4 mm or 7 mm. In the case of a block talc ring of × 7 mm × 3 mm (outer diameter × height × inner diameter in each case), the above-described loading variant is particularly suitable. Regarding the salt bath temperature, the statement of DE-A 443 19 49 applies. Usually selected in this way so that the acrolein conversion achieved in a single pass is usually 90 mol% or 95 mol% or 99 mol%.

However, the partial oxidation of acrolein to acrylic acid can also be carried out using the catalyst as described in DE-A 199 10 508, for example in a two-zone multi-catalyst tube fixed bed reactor. For the acrolein conversion, the above statements apply. Further, in the case where the partial oxidation of acrolein as described above is carried out as in the second reaction stage in which the two stages of propylene are partially oxidized to acrylic acid in a two-zone multi-catalyst tube fixed bed reactor, the gas mixture is filled (initial) The reaction gas mixture can suitably be used directly by using a partial gas oxidation of the product gas mixture for the first step (propylene → acrolein) (if appropriate, after indirect or direct (for example by supplying a second oxygen) intermediate cooling) Obtained (as described above). The oxygen required for partial oxidation of acrolein is preferably added in the form of pure molecular oxygen (if appropriate, also in the form of air or with molecular oxygen and preferably 50% by volume, better 40% by volume, or 30% by volume, or 20% by volume, even better 10% by volume, or 5 vol%, or 2 parts by volume of the inert gas fraction is added as a mixture of inert gases) and, for example, directly added to the product gas mixture of the first step of two-stage partial oxidation (propylene → acrolein). However, as already stated, it may also already be present in the starting reaction gas mixture of the first reaction stage.

Two-stage multi-catalyst tube fixed bed reactor (high reactants on the catalyst bed) in a two-stage partial oxidation of propylene to the acrylic acid in a second step of partially oxidizing the product gas mixture of the first step of partial oxidation further At a loading, such as a fairly normal condition, preferably a co-current mode between the reaction gas and the salt bath (hot carrier) observed via a tube bundle reactor) or two dual-zone multi-catalyst tube fixed bed reactors will typically Connect in series. Mixed series connections (single zone/dual zone or vice versa) are also possible.

An intercooler that can include, where appropriate, an inert bed that can perform the function of the filter can be placed between the reactors. The salt bath temperature of the multi-catalyst tube reactor for the first step of partial oxidation of propylene to acrylic acid is usually from 300 to 400 °C. For the second step of partial oxidation of propylene to acrylic acid, the salt bath temperature of the catalyst tube reactor in which the acrolein is partially oxidized to acrylic acid is usually from 200 to 350 °C. In addition, the heat exchange medium (preferably a salt melt) is usually characterized by the difference between its input temperature and its output temperature. An amount of 5 ° C was directed through the associated multi-catalyst tube fixed bed reactor. As already mentioned, the two steps of partial oxidation of propylene to acrylic acid can also be carried out as one part of the filler in one reactor as described in DE-A 101 21 592.

It should also be mentioned here that a part of the starting gas of the reaction gas B used in the first step ("propylene → acrolein") may be the residual gas from the separation zone II.

For the sake of completeness, it should be emphasized here that fresh propane can be additionally metered into the filling gas mixture of the first oxidation stage and the filling gas mixture of the second oxidation stage, or only one of the two mixtures. It is not preferred in accordance with the present invention, but it may be advantageous in some circumstances to eliminate the ignitability of the filling gas mixture.

In general, wherein the catalyst packing follows the first reaction step (the two-stage propylene partial oxidation in a single reactor (for example) by EP-A 911 313, EP-A 979 813, EP-A 990 636 and DE - A 28 30 765 teaches the completion of a tube bundle reactor which is suitably changed along individual catalyst tubes to form the simplest embodiment of the two oxidation stages for the two steps of partial oxidation of propylene to acrylic acid. If appropriate, the catalyst tube is filled with a catalyst and interrupted by a bed of inert material.

However, it is preferred to carry out the two oxidation stages in the form of two tube bundle systems connected in series. The systems can be placed in a reactor, in which case the conversion of one tube bundle to another is formed by a bed of inert material that is not supplied to the catalyst tubes (and suitably available in progress). Although the catalyst tube is typically surrounded by a hot carrier, it does not reach the bed of inert material supplied as described above. Advantageously, the two catalyst tube bundles are thus supplied to the spatially separated reactor. In general, an intercooler is disposed between the two tube bundle reactors to reduce any post-acrolein post combustion in the product gas mixture exiting the first oxidation zone. The reaction temperature in the first reaction stage (propylene → acrolein) is usually from 300 to 450 ° C, preferably from 320 to 390 ° C. The reaction temperature in the second reaction stage (acrolein → acrylic acid) is usually from 200 to 370 ° C, often from 220 to 330 ° C. The reaction pressure in the two oxidation zones is suitably from >1 to 4 bar, advantageously from 1.3 to 3 bar. The loading of the oxidation catalyst of the reaction gas in the two reaction stages (l(STP)/l.h) is often 1500 to 2500 l(STP)/l. h or to 4000 l(STP)/l. h. The loading of propylene or acrolein can be 100 to 200 or 300 and greater l (STP) / l. h.

In principle, the two oxidation stages in the method according to the invention can be configured as described in, for example, DE-A 198 37 517, DE-A 199 10 506, DE-A 199 10 508 and DE-A 198 37 519. .

In both reaction stages, the kinetics of partial oxygen partial oxidation with respect to the amount of molecular oxygen required in accordance with the amount of reaction chemistry required and having a beneficial effect on catalyst life.

In principle, it is also possible to carry out the heterogeneously catalyzed gas phase partial oxidation of propylene to acrylic acid in accordance with the invention in a single single zone tube bundle reactor as follows. The two reaction steps are carried out in an oxidation reactor packed with one or more catalysts, the active composition of which is a multi-metal oxide comprising the elements Mo, Fe and Bi and capable of catalyzing the reaction of two reaction steps . Of course, the catalyst charge can vary continuously or abruptly along the reaction coordinates. Of course, in the embodiment of the invention in which two stages of propylene are partially oxidized to acrylic acid in the form of two oxidation stages connected in series, it is possible (if desired) from the product gas before the product gas mixture is passed to the second oxidation stage. Part or all of the mixture is removed as carbon dioxide and steam formed as a by-product in the first oxidation stage and present in the product gas mixture leaving the first oxidation stage. Preferably, the removal procedure is not provided in accordance with the present invention.

As already mentioned, in addition to air (which is less preferred according to the invention), a suitable source for the intermediate oxygen feed between the two oxidation stages is pure molecular oxygen or with such as CO ₂ , CO Molecular oxygen diluted with an inert gas of a rare gas, N ₂ and/or a saturated hydrocarbon. Preferably, according to the invention, the source of oxygen comprises 50% by volume, preferably 40% by volume, better 30% by volume, even better 20% by volume, better 10% by volume, or 5 vol%, or 2% by volume of a gas other than molecular oxygen.

Suitable materials for the inert shaped dilutions in the catalyst bed for the two oxidation stages are those materials which have also been recommended for dilution of the catalyst bed in reaction zone A.

In the process according to the invention, the product gas mixture metered into, for example, the source of cold oxygen into the first partial oxidation stage may also be used before it is further used as a component of the starting reaction gas mixture of the second partial oxidation stage. A direct way to achieve its cooling.

The partial oxidation of acrolein to acrylic acid is advantageously carried out according to the invention as described in EP-A 1 159 246 and preferably as described in WO 04/085 365 and WO 04/085 370. However, it is preferred according to the invention to use the starting reaction gas mixture of the first stage propylene partial oxidation to acrolein product gas mixture of the present invention as the starting reaction gas mixture comprising acrolein, which is sufficient second when appropriate. The air is replenished so that the ratio of molecular oxygen to acrolein in the resulting initial reaction gas mixture is from 0.5 to 1.5 in each case. The documents EP-A 1 159 246, WO 04/08536 and WO 04/085 370 are considered as an integral part of this document.

In other words, partial oxidation of the acrolein to acrylic acid of the present invention can advantageously be carried out with an increased acrolein loading on a fixed catalyst bed having at least two temperature zones.

In general, the two-stage partial oxidation of propylene to acrylic acid is preferably carried out as described in EP-A 1 159 248 or WO 04/085 367 or WO 04/085 369.

The product gas B leaving the partial oxidation (after the first and/or second oxidation stage) to be carried out according to the invention substantially comprises acrolein or acrylic acid or a mixture thereof, unconverted propane, (where appropriate) molecules as target products Hydrogen, steam (formed as a by-product and/or additionally used as a diluent gas), any unconverted molecular oxygen (for the purpose of the life of the catalyst used, when the oxygen content of the product gas mixture in the two partial oxidation stages (eg ) is still at least 1.5 to 4 vol%, which is generally advantageous) and other by-products or minor components having a higher and lower boiling point than water (for example, CO, CO ₂ , low carbon aldehyde, lower hexane Carboxylic acids (such as acetic acid, formic acid and propionic acid), maleic anhydride, benzaldehyde, aromatic carboxylic acids and aromatic carboxylic anhydrides (such as phthalic anhydride and benzoic acid), in some cases other hydrocarbons, for example, C ₄ hydrocarbons (e.g. butene-1 and possibly other butenes), and other inert diluent gases in some cases, e.g. N _2).

Suitable methods for removing the target product present in the product gas B are in principle all methods known in the prior art in this respect, which are usually accompanied by the removal of water according to the method of the invention in the second separation zone II and have A secondary component that is higher than the boiling point of water. An essential feature of such processes is that the target product is, for example, converted from a gaseous state to a condensed phase by absorption and/or condensing measures, and the condensate or absorbent is subsequently extracted, distilled, crystallized and/or desorbed. Processed for the purpose of further target product removal. Together with the target product and/or subsequent conversion of the target product to a condensed phase, the processes also typically convert the vapor and the minor component present in the reaction gas B having a higher boiling point than water into a condensed phase, and thus Removal (in accordance with the present invention, preferably, the amount of steam converted to the condensed phase in the separation zone II is at least 70 mol%, preferably at least 80 mol%, more preferably at least 90% of the amount of steam formed in the reaction zone B. Mol% and more preferably at least 95 mol% (or more preferably all of the amount present in product gas B)).

Suitable absorbents include, for example, water, aqueous solutions and/or organic solvents (for example mixtures of biphenyl and diphenyl ether, or mixtures of biphenyl, diphenyl ether and dimethyl phthalate). This "condensation" (removal) of the target product, water, and secondary components having a higher boiling point than water typically leaves a residual gas that does not convert to a condensed phase and includes product gas B having a lower boiling point than water and A component that is quite difficult to condense. These components are usually (especially) whose boiling point is at standard pressure (1 bar) -30 ° C of their components (the total content in the residual gas is usually 60% by volume, often 70% by volume, in many cases 80% by volume, but usually 90% by volume). The components mainly comprise unconverted propane, carbon dioxide, any unconverted propylene, any unconverted molecular oxygen, (where appropriate) molecular hydrogen, and other minor components having a lower boiling point than water, such as CO, B. Alkanes, methane, in some cases N ₂ , in some cases rare gases (eg He, Ne, Ar, etc.). To a small extent, the residual gas may also include acrylic acid, acrolein and/or H ₂ O. Preferably, the residual gas comprises according to the invention 10% by volume, advantageously 5% by volume and particularly advantageous 3 vol% steam. The above residual gas typically forms (at least 80% or at least 90% or at least 95% or more) of the residual gas formed in the second separation zone II (based on the amount of propane present therein) and This is referred to herein as the (primary) residual gas.

Particularly when the target product is condensed by absorption by means of an organic solvent, at least a minor residual gas comprising unconverted propane and any unconverted propylene is usually obtained in the separation zone II (in terms of the propane present therein), The amount of (mainly) residual gas is usually substantially lower (usually 20%, usually 10% or 5%, or 1%)). This is due to the formation of a condensed phase that absorbs a certain amount of unconverted propane and any unconverted propylene.

In the further process of removing the target product from the condensed phase extraction, distillation, crystallization and/or desorption (in separation zone II), the unconverted propane and, where appropriate, propylene are generally used as at least one other gas phase group. The fraction is recovered and referred to herein as a (secondary) residual gas.

In this case, the sum of the (primary) residual gas and the (secondary) residual gas forms the total amount of residual gas remaining in the separation zone II. When no (secondary) residual gas is obtained in the separation zone II, the (primary) residual gas is automatically the total amount of residual gas (also referred to as (total) residual gas).

Preferably, according to the invention, the target product is converted from product gas B to a condensed phase by fractional condensation. This is especially true when the target product is acrylic acid. However, suitable methods for the removal of the target product are in principle all, for example, all of the absorption and/or condensation methods described in the following documents for further processing of "target product condensation" and "condensate": DE - A 102 13 998, DE-A 22 63 496, US 3, 433, 840 (the method described in the above references is especially recommended for acrolein removal), EP-A 1 388 533, EP-A 1 388 532, DE-A 102 35 847, EP-A 792 867, WO 98/01415, EP-A 1 015 411, EP-A 1 015 410, WO 99/50219, WO 00/53560, WO 02/09839, DE-A 102 35 847, WO 03/041833, DE-A 102 23 058, DE-A 102 43 625, DE-A 103 36 386, EP-A 854 129, US-A 4,317,926, DE-A 198 37 520, DE-A 196 06 877 , DE-A 190 50 1325, DE-A 102 47 240, DE-A 197 40 253, EP-A 695 736, EP-A 982 287, EP-A 1 041 062, EP-A 117 146, DE-A 43 08 087, DE-A 43 35 172, DE-A 44 36 243, DE-A 199 24 532, DE-A 103 32 758 and DE-A 199 24 533. Acrylic removal can also be carried out as described in EP-A 982 287, EP-A 982 289, DE-A 103 36 386, DE-A 101 15 277, DE-A 196 06 877, DE-A 197 40 252, DE-A 196 27 847, EP-A 920 408, EP-A 1 068 174, EP-A 1 066 239, EP-A 1 066 240, WO 00/53560, WO 00/53561, DE-A 100 53 086 and EP-A In 982 288. It is preferably removed as shown in Figure 7 of WO 0196271 or as described in DE-A 10 2004 032 129 and its equivalent. Advantageous methods of removal are also described in the documents WO 04/063138, WO 04/35514, DE-A 102 43 625 and DE-A 102 35 847. Further treatment of the crude acrylic acid obtained can be achieved, for example, as described in the documents WO 01/77056, WO 03/041832, WO 02/055469, WO 03/078378 and WO 03/041833.

Typically, product gas B is sent to separation zone II at the working pressure it leaves as it exits reaction zone B.

The common feature of the above separation method is (as already mentioned), mainly comprising the product gas B having a boiling point lower than the boiling point of water at a standard pressure (1 bar) and usually The residual gas stream of -30 ° C of their components (i.e., components that are difficult to condense or have a volatile component) is typically left on top of, for example, a particular separation column comprising separate internal components, and product gas B (for example) is typically It is sent to the lower section of the column after its direct and/or indirect cooling. However, the residual gas may, for example, still include components such as steam and acrylic acid.

In the lower section of the separation column, a relatively small volatile component of product gas B is typically obtained primarily in the condensed phase, containing a particular target product. The condensed aqueous phase is typically removed via the side draw and/or via the bottom.

Typically, the (primary) residual gas comprises the following contents: 1 to 20% by volume of H ₂ O, 0 to 80% by volume of N ₂ , 10 to 90% by volume of propane, 0 to 20% by volume of H ₂ , 0 to 10% by volume of O _2, 1 to 20% by volume of CO ₂ and 0 to 5% by volume of CO.

The content of acrolein and acrylic acid is usually <1% by volume in each case. Preferred (primary) residual gases include: 1 to 5% by volume of H ₂ O, 10 to 80% by volume of propane, 0 to 5% by volume of N ₂ , 0 to 20% by volume of H ₂ , and 1 to 10% by volume. O ₂ , 0 to 5 vol% of CO and 1 to 15 vol% of CO ₂ .

Generally, in the method according to the invention, the residual gas is 1 bar and 3 bar, preferably 2.5 bar and usually The pressure of 2 bar leaves the separation zone II. The temperature of the residual gas is usually 100 ° C, usually 50 ° C (but usually 0 ° C).

Suitably, the propane present in the residual gas is suitably absorbed in the separation zone III at a pressure of from 5 to 50 bar, preferably from 10 to 25 bar, from the viewpoint of application. In the compression zone, the residual gas leaving the separation zone II will typically be compressed to a pressure that facilitates the absorption removal of propane (which is advantageously achieved in several stages). Preferably, according to the invention, the compression of the residual gas is in the first compression stage by means of a multistage (for example) turbocompressor (radial compressor; for example, MH4B from Mannesmann DEMAG, Germany) (here also The residual gas compressor is carried out to a working pressure substantially corresponding to the pressure in the reaction zone A (for example, 1.20 bar to 2.40 bar; in general, the compression system is achieved in the first compression stage to be slightly larger than the reaction zone A The pressure of the work pressure used in it). This heats the residual gas, for example, up to 50 °C. When not all of the residual gas is absorbed by propane, the compressed residual gas as described is appropriately divided into two portions having the same composition. The proportion of the smaller portion may, for example, be up to 49% of the total amount, preferably no more than up to 40%, or up to 30%, or up to 20%, or up to 10%. It can also be the total amount 10% and 49%.

This portion can be used as a residual gas recycle gas for immediate recycle to reaction zone A. The other part is further compressed to the working pressure expected for propane absorption. It can in principle be implemented only in another compression phase. However, advantageously according to the invention at least two further compression stages will be used for this purpose. This is due to the compression of the gas (as noted in the first compression stage) associated with an increase in temperature. Conversely, its associated cooling "of propane washed" under reduced pressure of the residue gas (expansion) (which have appropriate H ₂ film removal). When the reduced pressure is also carried out in several stages, the indirect heat exchange between the heated, compressed, unwashed residual gas and the cooled, expanded, "propane-washed" residual gas can be used under each condition. Before the stage. The heat exchange (which may additionally be supported by air cooling between the compressed residual gases) may also be used prior to the first expansion (due to the indirect heat exchange, condensation of any steam still present in the residual gas may occur therein) . Advantageously, from the application point of view, the expansion described above is carried out in an (advantageously multi-stage) expansion turbine (which is used for the recovery of the compression energy; in other words, the mechanical energy obtained during the expansion can be used in all cases (especially in the examples and In the comparative example) additional or primary drive and/or electricity generation directly used as one of the compressors. Depending on the compression or expansion stage, the difference between the inlet and outlet pressures can be, for example, 2 to 15 bar.

The compressed residual gas may be before the propane present in the compressed residual gas is in the separation zone III from the compressed residual gas absorption inhalation organic solvent (the compressed residual gas is washed by means of an organic solvent (absorbent)) Separating through another membrane in another separation zone to remove at least a portion of any molecular hydrogen still present therein. Suitable separation membranes in this regard are, for example, aromatic polyimine membranes, such as those from UBE Industries Ltd. Among the latter, film types A, B-H, C and D are particularly suitable. For this removal, it is especially preferred to use a UBE Industries Ltd. B-H polyimide film. At 60 ° C, the H ₂ hydrogen permeation rate of the film is 0.7.10 ^-3 [STP. Cc/cm ² . Sec. Cm Hg]. For this purpose, the compressed residual gas can pass through a membrane that generally has a tubular shape (but a plate or roll module is also suitable) and is only permeable to molecular hydrogen. The removed molecular hydrogen can therefore be further used for other chemical synthesis or at least partially recycled to reaction zone A (suitably as component of the reaction gas A charge gas mixture).

According to the invention, it is suitable to separate molecular hydrogen from the compressed residual gas, since it is preferably also carried out under high pressure (for example 5 to 50 bar, usually 10 to 25 bar). Advantageously, according to the invention, the propane absorption of the residual gas (propane wash) is carried out immediately or alternatively (and preferably according to the invention) after the propane wash of the residual gas (propane absorption) is carried out immediately (as appropriate) The hydrogen membrane separation from the residual gas is performed to double the pressure level that has been advantageously increased once. In addition to the plate-like film, a roll film or capillary film, a tubular film (hollow fiber membrane) is particularly suitable for hydrogen removal. The inner diameter can be, for example, from several micrometers to several micrometers. For this purpose, a tube similar to a tube bundle reactor, for example, a bundle of such tubular membranes, in each case, is cast into a plate at each end of the tube. Outside the tubular membrane, it is preferred to have a reduced pressure (<1 bar). At high pressure, residual gas (or "propane-washed" residual gas) is directed toward one of the two plate ends and forced through the inside of the tube to the tube outlet present at the end of the opposite plate. Along the interior of the tube thereby defining the flow path, molecular hydrogen is released outward through the H ₂ - permeable membrane.

The absorption of propane from the residual residual gas (the unconverted propylene which is partially oxidized during the partial oxidation in reaction zone B is generally absorbed together with propane) is substantially not limited at all. A simple way for this purpose, for example, is to leave residual residual gas at a pressure of from 5 to 50 bar, preferably from 10 to 25 bar (according to the invention suitably, it has from 10 to 100 ° C, preferably from 10 to The temperature of 70 ° C is in contact with an organic solvent (preferably hydrophobic), and the temperature of the organic solvent is advantageously from 0 to 100 ° C, preferably from 20 to 50 ° C, or to 40 ° C, and wherein propane and propylene are absorbed. (appropriately better than the other components of the remaining residual gas) (for example by simply passing it).

For example, subsequent desorption (rapid evaporation), rectification, and/or use is inert (eg, N ₂ ) relative to reaction zone A and/or required as a reactant (eg, air or molecular oxygen) in the reaction zone Stripping of the gas with another mixture of inert gas, preferably steam or a mixture of steam and molecular oxygen, or a mixture of steam and air, may remove propane as a mixture from the absorbent in separation zone IV (suitable Propylene, which can be recycled to reaction zone A as a feed stream comprising propane.

Suitable absorbents for the above mentioned absorbent removal are in principle all absorbents which are capable of absorbing propane and propylene. The absorbent is preferably a hydrophobic and/or high boiling organic solvent. Advantageously, the solvent has (at a standard pressure of 1 bar) a boiling point of at least 120 ° C, preferably at least 180 ° C, preferably from 200 to 350 ° C, especially from 250 to 300 ° C, more preferably from 260 to 290 ° C. Suitably, the flash point (at a standard pressure of 1 bar) is greater than 110 °C. Suitable as absorbents are generally relatively non-polar organic solvents, such as aliphatic hydrocarbons which preferably do not include any externally reactive polar groups, but are also aromatic hydrocarbons. Generally, it is desirable that the absorbent has an extremely high boiling point while having extremely high solubility to propane and propylene. Examples of the absorbent include aliphatic hydrocarbons such as C ₈ -C ₂₀ alkanes or alkenes, or aromatic hydrocarbons such as those derived from a paraffinic distillation intermediate oil fraction or having a bulky (space-required) group on the oxygen atom. Or a mixture thereof, a polar solvent such as dimethyl 1,2-phthalate disclosed in DE-A 43 08 087 may be added thereto. Also suitable are esters of benzoic acid and phthalic acid with linear alkanols having from 1 to 8 carbon atoms, such as n-butyl benzoate, methyl benzoate, ethyl benzoate, dimethyl phthalate Esters, diethyl phthalate, and those known as hot carrier oils, such as biphenyl, diphenyl ether and mixtures of biphenyl and diphenyl ether or their chloro derivatives and triarylalkanes, such as 4-methyl -4'-Benzyldiphenylmethane and its isomers, 2-methyl-2'-benzyldiphenylmethane, 2-methyl-4'-benzyldiphenylmethane and 4-methyl -2'-Benzyldiphenylmethane and a mixture of such isomers. Suitable absorbents are solvent mixtures of biphenyl and diphenyl ether, preferably in an azeotropic composition, especially a solvent mixture of about 25% by weight of biphenyl and about 75% by weight of diphenyl ether, such as commercially available Diphyl. (eg from Bayer Aktiengesellschaft). Often, the solvent mixture includes a solvent such as dimethyl phthalate added in an amount of 0.1 to 25% by weight based on the entire solvent mixture. Particularly suitable absorbents are also octane, decane, decane, undecane, dodecane, tridecane, tetradecane, pentadecane, hexadecane, heptadecane and octadecane, especially Tetradecane has been found to be particularly suitable. It is advantageous when the absorbent used first satisfies the above boiling point and secondly does not have an excessively high molecular weight. Advantageously, the molecular weight of the absorbent is 300 g/mol. Also suitable are paraffinic oils having from 8 to 16 carbon atoms as described in DE-A 33 13 573. Examples of suitable products are those sold by Haltermann, such as Halpasols i, such as Halpasol 250/340 i and Halpasol 250/275 i, and printing inks designated PKWF and Printosol. Preferred are aromatics-free commodities, such as those having the PKWFaf type. If it contains a small amount of residual aromatics, it can be advantageously reduced and reduced to a value of substantially less than 1000 ppm by weight prior to use, advantageously by rectification and/or absorption. Other suitable commodities are paraffin wax (C ₁₃ -C ₁₇ ) or from Erd Mihagol of l-Raffinerie-Emsland GmbH 5. LINPAR from CONDEA Augusta SpA (Italy) or SASOL Italy SpA 14-17. Paraffin wax (heavy) C ₁₄ -C ₁₈ from SLOVNAFT of Slovakia.

The contents of the above products with linear hydrocarbons (reported by area analysis by gas chromatography) are generally: total C ₉ to C ₁₃ : less than 1%; C ₁₄ : 30 to 40%; C ₁₅ : 20 to 33%; C _16: 18 to 26%; C _17: up to 18%; :<2%.

Typical compositions of products from SASOL are: C ₁₃ : 0.48%; C ₁₄ : 39.8%; C ₁₅ : 20.8%; C ₁₆ : 18.9%; C ₁₇ : 17.3%; C ₁₈ : 0.91%; C ₁₉ : 0.21% .

The typical composition of the product from Haltermann is: C ₁₃ : 0.58%; C ₁₄ : 33.4%; C ₁₅ : 32.8%; C ₁₆ : 25.5%; C ₁₇ : 6.8%; : <0.2%.

In continuous operation, the composition of the absorbent will vary accordingly depending on the method.

The absorption can be carried out in a column or quench unit. It may be operated in a cocurrent or (preferably) countercurrent manner. The absorption column is adapted (e.g.) a tray column (having blisters and / or sieve trays), having a structured packing (e.g., of 100 to 1000 m ² / m ³ or to 750 m ² / m ³ a specific surface area of the sheet Metal fillers such as Mellapak 250 Y) tower and tower with random packing (eg filled with Raschig packing). However, it is also possible to use spray towers and spray towers, graphite block absorbers, surface absorbers such as thick film and film absorbers, as well as plate scrubbers, cross spray scrubbers and rotary scrubbers. Furthermore, it can be advantageously achieved that the absorption takes place in a bubble column with or without internal components.

Propane and, where appropriate, propylene can be removed from the absorbent (self-absorbent) by stripping, flash evaporation (flashing) and/or distillation (rectification).

Propane and propylene are preferably removed from the absorbent by vapor mention and/or desorption. Desorption can be carried out in a customary manner by means of pressure and/or temperature changes, for example at a pressure of from 0.1 to 10 bar, or from 1 to 5 bar, or from 1 to 3 bar, and from 0 to 200 ° C, especially from 20 to 100 ° C. More preferably, it is carried out at a temperature of 30 to 70 ° C, particularly preferably 30 to 50 ° C. The stripping gas used for stripping is preferably steam. A suitable alternative is molecular nitrogen or a mixture of molecular nitrogen and steam. However, carbon dioxide or a mixture thereof with steam and/or nitrogen is also a stripping gas suitable in accordance with the present invention. Particularly suitable for stripping is also a bubble column with and without internal components.

Propane and, where appropriate, propylene may also be removed by means of distillation or rectification from the absorbent (self-absorbent), in which case it is familiar to those skilled in the art and having structured packing, random packing or corresponding internal components. Tower. Preferred conditions in the distillation or rectification are from 0.01 to 5 bar, or from 0.1 to 4 bar, or from 1 to 3 bar, and from 50 to 300 ° C, especially from 150 to 250 ° C (at the bottom).

The propane recycle gas obtained by stripping the self-absorbent may be sent to another processing stage in the separation zone IV before it is used to charge the reaction zone A, for example to reduce the loss of the carry-over absorbent (eg, in addition to Separation in the water and/or deep filter) and thus simultaneously protect the reaction zone from the absorbent. This removal of the absorbent can be accomplished by all of the method variants known to those skilled in the art. An example of a preferred embodiment of the removal in accordance with the method of the present invention is to quench the outlet stream from the stripping unit with water. In this case, the absorbent is washed out of the filled outlet stream with water while simultaneously carrying the outlet stream with water. This washing, i.e., quenching, can be accomplished, for example, at the top of the desorption column by means of a liquid collection tray, by reverse spray of water or in a dedicated apparatus. To facilitate the separation effect, it is possible to install internal components known to those skilled in the art that self-rectify, absorb, and desorb the known quenching surface area in the quench chamber.

The pressure reduction of the absorbent comprising propane (and propylene if appropriate) to a pressure suitable for the removal of propane (and propylene if appropriate) can be carried out in the process according to the invention, for example by means of a valve or by means of a reverse pump. The mechanical energy released in the case of a reverse pump is suitably also used according to the invention to recompress an absorbent containing no propane (and, if appropriate, propylene) (for example in a desorption column or in a stripping column). Typically, the propane-free absorbent can be reused as an absorbent (according to the invention, suitably re-compressed to absorption pressure by means of a pump) to remove propane from the residual gas in separation zone III.

Particularly advantageously according to the invention, the propane is self-absorbed in the separation zone IV at a pressure of 1.5 or 2 to 5 bar, preferably > 2 to 5 bar, more preferably 2.5 to 4.5 bar and most preferably 3 to 4 bar. Object removal.

Since the propane recycle gas obtained at these working pressures can be directly recycled to the reaction zone A, it is extremely advantageous to remove the propane from the absorbent at the above working pressure. The propane recycle gas can be used as an actuating spray for a dehydrogenation recycle gas stream by the jet pump principle according to the teaching of DE-A 102 11 275. Suitably, according to the invention, the propane recycle gas system is suitably heated according to the invention to a temperature suitable for the reaction zone A by heat exchange with the hot reaction gas A and/or with the hot product gas A in advance.

A typical inclusion of a propane recycle gas obtained by desorption and/or stripping of the vapor comprising the propane by means of steam in the separation zone IV can be as follows in the process according to the invention: 80 to 99.99 Mol% propane, 0 to 5 mol% propylene, and 0 to 20 mol% H ₂ O. In many cases, the contents of the propane recycle gas are: 80 to 99 mol% propane, 1 to 4 mol% propylene, and 2 to 15 mol% H ₂ O.

In addition to fresh propane (as a component of crude propane) and propane recycle gas and, where appropriate, dehydrogenation recycle gas, (with appropriate compression) residual gas recycle gas and / or one (other) gas including molecular oxygen In the case of being sent to the reaction zone A, the crude propane and the residual gas recycle gas and/or the gas including molecular oxygen (appropriately from the application point of view) are premixed to obtain the first gas mixture and the first gas mixture (preferably in It has been mixed with a mixture of a propane recycle gas or a dehydrogenation recycle gas by heat exchange with the hot reaction gas A and/or the product gas A (preferably also by reacting with the hot reaction gas A and / or the product gas A is exchanged for heat exchange and correspondingly heated) and the resulting gas mixture is introduced into the reaction zone A as a reaction gas A filling gas mixture. Advantageously, molecular hydrogen is also added to the resulting gas mixture, after which the gas mixture is fed as a reactive gas A to the reaction zone A.

In the process according to the invention, the demand for a gas comprising molecular oxygen (appropriate from an application point of view) is taken from a common source (contained by ordinary sources).

According to the invention, the crude propane is suitably evaporated by means of indirect heat exchange by means of the condensate obtained in the separation zone II (for example acid water) (the thus cooled condensate can be used for direct cooling in the separation zone II to obtain New condensate). Alternatively, other liquid phases (which suitably have a temperature of 20 to 40 ° C, preferably 25 to 35 ° C) can also be used as a heat carrier for the exchange of heat between the purpose of evaporating the crude propane. In order to further heat the crude propane, it is usual to use a heat exchange with the hot steam (which is available, for example, at 5 bar and 152 ° C), which is usually used as a method according to the invention (for example, in removal) A by-product of the process of heat treatment obtained during partial oxidation is obtained in large quantities.

If desired, a small amount of absorbent is continuously unloaded from the separation zone IV and replaced by a fresh absorbent. The unloaded absorbent can be treated by rectification to obtain a fresh absorbent. The residual gas which has been subjected to reduced pressure, previously "washed with propane" and then (if appropriate) removed by membrane hydrogen is usually sent to the residue for incineration (exhaust gas incineration). The aqueous condensate (for example, acid water) formed between the high-boiling substance removed in the separation zone II and the separation zone II and, where appropriate, between the separation zone II and the separation zone III may also be introduced into the incineration (the incineration) When appropriate, it is supported by feeding natural gas to support and using air as a source of combustion oxygen. The incineration hot exhaust gas is typically used to exchange heat exchange with water to produce steam and is typically subsequently released to the atmosphere. Typically, it includes only molecular oxygen, molecular nitrogen, molecular hydrogen, and carbon dioxide. Finally, it should be emphasized that the same polymerization inhibitor is added whenever the condensation phase comprising acrylic acid and/or acrolein is present in the process according to the invention. Suitable polymeric inhibitors are in principle all known treatment inhibitors. Particularly suitable according to the invention are, for example, phenothiazine and methyl ether of hydroquinone. The presence of molecular oxygen increases the effectiveness of the polymerization inhibitor.

The acrolein obtained by the process according to the invention can be converted into the acrolein subsequent product mentioned in the documents US-A 6,166,263 and US-A 6,187,963. These products comprise 1,3-propanediol, methionine, glutaraldehyde and 3-methylpyridine.

Preferred is the method according to the invention having the maximum working pressure in the separation zone III.

Advantageously, for the working pressure P in different zones of the method according to the invention (measured at the entrance to a particular zone in each case), the following correlation (relationship) applies: P _{reaction zone A} > P _{separation Zone I} > P _{reaction zone B} > P _{separation zone II} < P _{separation zone III} > P _{separation zone IV} > P _{reaction zone A.}

Although the separation zone I can also be omitted.

It is also advantageous for the process according to the invention to maintain the temperature of the reaction gas A (especially when it comprises molecular oxygen) outside the catalyst and/or inert bed. 460 ° C, preferably A value of 440 ° C in order to thus minimize improper combustion reactions and thermal decomposition, in particular the combustion reaction and thermal decomposition of propane.

Examples and comparison examples: I. Long-term operation of propylene heterogeneous phase catalysis for two-stage partial oxidation to acrylic acid in the absence and presence of hydrogen molecules

A) General Experiment of Reaction Apparatus A reactor for the first oxidation stage is set. The reactor consists of a jacketed cylinder made of stainless steel (a cylindrical conduit surrounded by a cylindrical outer vessel). The wall thickness is often 2 to 5 mm.

The outer cylinder has an inner diameter of 91 mm. The inner diameter of the catheter is approximately 60 mm.

At the top and bottom, the jacket cylinders are closed by a lid and a bottom, respectively.

The catalyst tube (total length 400 cm, inner diameter 26 mm, outer diameter 30 mm, wall thickness 2 mm, stainless steel) is housed in a conduit of a cylindrical container so that it protrudes through the cover at its upper and lower ends in each condition And the bottom (in a sealed manner). The heat exchange medium (salt melt composed of 53% by weight of potassium nitrate, 40% by weight of sodium nitrite and 7% by weight of sodium nitrate) was enclosed in a cylindrical vessel. To ensure a very uniform thermal boundary condition of the outer wall of the catalyst tube over the entire length (400 cm) of the catalyst tube in the cylindrical vessel, the heat exchange medium is pumped in a circular manner by means of a rotary vane pump.

An electric heater connected to the outer jacket adjusts the temperature of the heat exchange medium to a desired level. In addition, there is air cooling.

Reactor Filler: Viewed through the first stage reactor, the salt melt and the charge gas mixture of the first stage reactor are directed in a cocurrent flow. The filling gas mixture enters the first stage reactor at the bottom. In each case, it was introduced into the reaction tube at a temperature of 165 °C. The salt melt enters the cylindrical conduit at the bottom at a temperature T ⁱⁿ and exits the cylindrical conduit at the top at a temperature ^Tout above T ⁱⁿ 2 °C. The T ^{in is} adjusted so as to produce a propylene conversion of 97.8 ± 0.1 mol% at the outlet of the first oxidation stage in a single pass.

Catalyst tube packing: (from the bottom up) Section A: The upstream bed of a talc ball of length 90 cm in diameter and 4-5 mm. Section B: Length 100 cm 30% by weight Geometry 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter) block talc ring and 70% by weight of the unsupported catalyst from the C section Catalyst filler for the mixture. Section C: 200 cm in length according to Example 1 of DE-A 100 46 957 (5 mm × 3 mm × 2 mm = outer diameter × length × inner diameter) catalyst loading of unsupported catalyst (stoichiometry: [Bi ₂ W ₂ O ₉ x2WO ₃ ] _0.5 [Mo ₁₂ Co _5.5 Fe _2.94 Si _1.59 K _0.08 O _x ] ₁ ). Section D: length 10 cm geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) of the bed downstream of the talc ring

Intermediate cooling and intermediate oxygen feed (pure O ₂ as second gas) for the purpose of intercooling (indirectly by means of air), the product gas mixture leaving the first fixed bed reactor is guided through the connecting pipe (length 40 cm, Inner diameter 26 mm, outer diameter 30 mm, wall thickness 2 cm, stainless steel, wrapped by 1 cm of insulating material), the connecting tube is mounted at the center to a length of 20 cm, geometry 7 mm × 3 mm × 4 mm ( The outer bed of the talc ring of outer diameter x length x inner diameter is filled and directly flanged on the first stage catalyst tube.

Total product gas mixture into the connecting tube (first stage) in> T ⁱⁿ the temperature is higher than 200 ℃ and at a temperature below 270 deg.] C and away from the connector tube. At the end of the connecting tube, molecular oxygen is metered into the cooled product gas mixture at the pressure level of the product gas mixture. The resulting gas mixture (the filling gas mixture of the second oxidation stage) is introduced directly into the second stage catalyst tube, which is also flanged to the catalyst tube by the other end thereof. The amount of molecular oxygen charged is such that the molar ratio of O ₂ present in the resulting gas mixture to acrolein present in the resulting gas mixture is 1.3.

The reactor of the second oxidation stage uses a catalyst tube fixed bed reactor having the same design as the one used for the first oxidation stage. Viewed through the reactor, the salt melt and the fill gas mixture were directed in a cocurrent flow. The salt melt enters at the bottom and the filling gas mixture also enters at the bottom. Adjusting the salt melt inlet temperature T ⁱⁿ order to produce a total of 99.3 ± 0.1 mol% acrolein conversion at the outlet of the second oxidation stage in a single pass. The T ^{out of the} salt melt is higher than T ⁱⁿ by 2 °C.

The catalyst tube packing (from the bottom up) is: A section: length 70 cm geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) of the preparation bed of the block talc ring.

Section B: length 100 cm 30% by weight geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) block talc ring and 70% by weight of coated catalyst from section C uniform Catalyst filler for the mixture.

Section C: Length of 200 cm According to DE-A 10046928 Preparation of Example 5 Ring (7 mm × 3 mm × 4 mm = outer diameter × length × inner diameter) coated catalyst catalyst filler (stoichiometry: Mo ₁₂ V ₃ W _1.2 Cu _2.4 O _x ).

Section D: The downstream bed of a talc ball of 30 cm in diameter and 4-5 mm in diameter.

B) the result of a function as a function of the composition of the filling gas mixture in the first oxidation stage (propylene loading is set to 150 l (STP) / l.h; the selectivity for acrylic acid formation is always 94 mol%).

a) The composition of the filling gas mixture in the first oxidation stage is substantially: 7% by volume of propylene, 12% by volume of O ₂ , 20% by volume of H ₂ , 5% by volume of H ₂ O and 56% by volume of N ₂ .

At the start of the reaction apparatus freshly packed with the catalyst, the inlet temperature was: T ⁱⁿ (first oxidation stage): 320 °C.

T ⁱⁿ (second oxidation stage): 268 ° C.

After 3 months of operation time, the inlet temperature was: T ⁱⁿ (first oxidation stage): 330 °C.

T ⁱⁿ (second oxidation stage): 285 ° C.

b) the composition of the gas mixture of the first oxidation stage of filling substantially: 7% by volume of propylene, 12 vol% of O _2, 5% by volume of H ₂ O and 76% by volume of N _2.

At the beginning of the reaction apparatus freshly packed with the catalyst, the inlet temperature was: T ⁱⁿ (first oxidation stage): 320 °C.

T ⁱⁿ (second oxidation stage): 268 ° C.

After 3 months of operation time, the inlet temperature was: T ⁱⁿ (first oxidation stage): 324 °C.

T ⁱⁿ (second oxidation stage): 276 °C.

II. First illustrative method for preparing acrylic acid from propane (describes stable operating conditions)

Reaction zone A consists of a blast furnace reactor which has been designed as a tray reactor and which is arranged in an adiabatic configuration and has three fixed catalyst beds arranged in sequence in the flow direction. The specific fixed catalyst bed is placed on a stainless steel wire mesh with inert materials (arranged in a predetermined order in the flow direction) (bed height: 26 mm; talc balls 1.5 to 2.5 mm in diameter) and dehydrogenation catalyst: block talc Ball bed volume ratio = 1:3 bed of a mixture of fresh dehydrogenation catalyst and talc balls (1.5 to 2.5 mm in diameter) (or, nevertheless, it is also possible to use the same amount of undiluted catalyst here).

A static gas mixer is placed upstream of each fixed bed. The dehydrogenation catalyst has been promoted with the elements Cs, K and La in oxidized form and has been used for ZrO ₂ . Outer surface and inner surface of SiO ₂ mixed oxide carrier extrudate (average length (Gaussian distribution in the range of 3 mm to 12 mm, maximum of about 6 mm): 6 mm, diameter: 2 mm) The Pt/Sn alloy whose elemental stoichiometry is (including the mass ratio of the carrier) Pt _0.3 Sn _0.6 La _3.0 Cs _0.5 K _0.2 (ZrO ₂ ) _88.3 (SiO ₂ ) _7.1 (as in the catalyst of Example 4 of DE-A 10219879) The precursor is prepared and activated into an active catalyst).

The heterogeneously catalyzed partial propane dehydrogenation is carried out in the tray reactor in a direct transfer mode with a dehydrogenation recycle gas. The total amount of catalyst in all disks versus propane loading (calculated without inert material) is 1500 l(STP)/l. h.

7.30 m ³ (STP) / h of the reaction gas A filling gas mixture (starting mixture: P = 2.13 bar) is sent to the first catalyst bed in the flow direction and has the following contents: propane 61.22 vol%, propylene 12.17 vol. %, ethylene 0.86 vol%, methane 2.56 vol%, H ₂ 3.78 vol%, O ₂ 2.10 vol%, H ₂ O 12.08 vol%, CO 0.38 vol%, and CO ₂ 3.35 vol%.

A reaction gas mixture of the gas mixture filling the gas mixture and the second ^{(4.96 m 3 (STP) /} h) first ((STP) / h 2.34 m 3), by means of a static mixer which is based a combination of two gas Obtained from the mixture. The first gas mixture has a temperature of 400 ° C and a pressure of 2.13 bar and comprises: 1.19 m ³ (STP) / h of compressed residual gas recycle gas (T = 70 ° C, P = 2.40 bar), 1.05 m ³ (STP) / a crude propane of h comprising propane (fresh propane) to a level of >98 vol% (T = 237 ° C, P = 2.90 bar), and O ₂ (T = 139 ° C, 0.098 m ³ (STP) / h, P = 2.33 bar, purity > 99% by volume).

The indirect heat exchange with the product gas A in the first heat exchanger with the derivation reaction zone A and in the direction of the reaction zone B produces a temperature of 400 °C. The second gas mixture comprises: 3.68 m ³ (STP) / h of dehydrogenation cycle gas (T = 551 ° C, P = 2.05 bar) and 1.28 m ³ (STP) / h of propane cycle gas (T = 400 ° C, P = 3.26 bar).

It has a pressure of 2.13 bar and a temperature of 506 °C.

Indirect heat exchange with product gas A cooled to a temperature of 435 ° C in the first heat exchanger produces a temperature of the propane recycle gas. The dehydrogenation cycle gas stream is achieved according to the principle of a jet pump (cf. DE-A 10211275), and the propane cycle gas heated to 400 ° C acts as an actuating jet.

Adjusting the inlet temperature of the reaction gas A charging gas mixture and the bed height of the first catalyst bed through which the reaction gas A is filled with the gas mixture so that the reaction gas A having the following contents leaves at a temperature of 515 ° C and a pressure of 2.11 bar Catalyst bed: propane 55.8 vol%, propylene 15.85 vol%, ethylene 0.88 vol%, methane 2.89% by volume, H ₂ 3.78 vol%, O ₂ vol%, H ₂ O 15.5 vol%, and CO ₂ 3.67 vol%.

The amount of departure is 7.46 m ³ (STP) / h.

When the first catalyst bed was passed, 0.12 m ³ (STP) / h of molecular oxygen (purity > 99 vol%) was metered into the reaction gas A (T = 139 ° C, P = 2.33 bar). The metered addition was achieved by throttling so that the resulting pressure of the resulting reaction gas A was still 2.11 bar.

The bed height of the second catalyst bed was such that the reaction gas A having the following contents left the second catalyst bed at a temperature of 533 ° C and a pressure of 2.08 bar: propane 49.2% by volume, propylene 19.24% by volume, ethylene 1.04% by volume, methane 3.14% by volume, H ₂ 4.43 vol%, O ₂ 0 vol%, H ₂ O 17.98 vol%, and CO ₂ 3.52 vol%.

The amount of departure is 7.79 m ³ (STP) / h.

Disposed upstream of a static mixer upstream of the third catalyst bed, molecular 0.13 m ³ (STP) / h of oxygen (purity> 99% by volume) the amount of the reactive gas A (T = 139 ℃, 2.33 Throttling under the pressure of Ba). The resulting pressure of the obtained reaction gas A was still 2.08 bar.

The bed height of the third catalyst bed was such that the reaction gas A exited the third catalyst bed as a product gas A having the following contents at a temperature of 551 ° C and a pressure of 2.05 bar: propane 42.19 vol%, propylene 22.79 vol%, ethylene 1.14. % by volume, methane 3.42% by volume, H ₂ 5.28% by volume, O ₂ 0% by volume, H ₂ O 20.32% by volume, and CO ₂ 3.36% by volume.

The amount of departure is 8.17 m ³ (STP) / h.

The product gas A is divided into two parts having the same composition by a flow divider. The first part is 3.68 m ³ (STP)/h and the second part is 4.49 m ³ (STP)/h. The first portion is recycled to the reaction zone A as a component of the reaction gas A charging gas mixture. For the purpose of the first part of the dehydrogenation recycle gas stream, the heated propane recycle gas is used as an actuating injection to operate the jet pump, and the transport direction of the decelerated actuated jet passes through the mixing zone through the actuating nozzle, And the diffuser is directed in the direction of the inlet of the reaction zone A, and the nozzle is directed in the direction of the first flow portion of the product gas A, and the "nozzle-mixing zone-diffuser" is connected to form the first part of the product gas A to be recycled The only connecting line between the picking up of reaction zone A. The second portion of product gas A is directed to reaction zone A and first cooled from 551 ° C to 435 ° C by a first indirect heat exchange with a first gas mixture comprising a compressed residual gas recycle gas, crude propane And molecular oxygen, contained in the starting mixture of reaction gas A (filling gas mixture). When it is completed, the working pressure of product gas A deriving reaction zone A drops from 2.05 bar to 1.95 bar. At the same time, the first gas mixture was heated from 164 ° C to 506 ° C and reduced from 2.40 bar to 2.13 bar.

The product gas A was cooled from 435 ° C to 335 ° C in a second indirect heat exchange downstream from the propane recycle gas from separation zone IV. When it is completed, the working pressure in product gas A drops to 1.94 bar. Conversely, the temperature of the propane recycle gas increased from 69 ° C to 400 ° C.

Subsequently, 0.76 m ³ (STP) / h of steam was removed by condensation from product gas A cooled to 335 ° C. To this end, the product gas A, which is cooled to 335 ° C, is indirectly indirectly with an appropriate amount of steam which has been removed by condensation to a product gas A ^* (T = 40 ° C, P = 1.68 bar) of 3.73 m ³ (STP) / h. Heat exchange, and therefore cooled to 108 °C. The temperature of the product gas A ^* was simultaneously increased to 305 °C. In the downstream air cooler, product gas A (T = 108 ° C, P = 1.94 bar) is cooled to 54 ° C by heat exchange with external air (the working pressure of product gas A drops to 1.73 bar) and thereafter Has been biphasicity.

Direct cooling with atomized aqueous condensate which has been cooled (to 29 ° C) and previously removed from product gas A by suitable direct cooling can be obtained by condensing the above amount of steam from product gas A to obtain product gas A ^* (P = 1.68 bar, T = 40 ° C, 3.73 m ³ (STP) / h). It was heated to 305 ° C as described above. Molecular oxygen (purity > 99% by volume, T = 139 ° C, P = 2.33 bar) of 1.39 m ³ (STP) / h was then throttling into the product gas A ^* . It forms a reaction gas B starting mixture (filling gas mixture) having a temperature of 288 ° C and a working pressure of 1.63 bar.

The reaction gas B starting mixture (filling gas mixture) had the following contents: propane 37.01 vol%, propylene 19.99 vol%, ethylene 0.99 vol%, methane 2.99% by volume, H ₂ 4.63 vol%, O ₂ 26.79 vol%, H ₂ O 3.14 vol%, and CO ₂ 2.95 vol%.

It is used to fill reaction zone B.

The first oxidation stage in the flow direction of the reaction gas B is a multi-tube reactor having two temperature zones. The reaction piping is configured as follows: V2A steel; outer diameter 30 mm, wall thickness 2 mm, inner diameter 26 mm, length 350 cm. From the top down, the reaction tube was packed as follows: 1 section: 50 cm in length as a block talc ring with an upstream bed geometry of 7 mm x 7 mm x 4 mm (outer diameter x length x inner diameter).

2 section: length 140 cm 20% by weight (or 30% by weight) geometry 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter) block talc ring and 80% by weight (or 70% by weight) A catalyst mixture from a homogeneous mixture of 3-stage unsupported catalysts.

Section 3: Length 160 cm The catalyst packing of the unsupported catalyst according to Example 1 of DE-A 100 46 957 (5 mm × 3 mm × 2 mm = outer diameter × length × inner diameter) (stoichiometry: [Bi ₂ W ₂ O ₉ .2WO ₃ ] _0.5 [Mo ₁₂ Co _5.5 Fe _2.94 Si _1.59 K _0.08 O _x ] ₁ ). Alternatively, it is also possible to use one of the catalysts EUC 1 to EUC 11 from Research Disclosure No. 497012 of August 29, 2005 (wherein the specific surface area of the active composition is erroneously described in cm ² /g; however The correct size is m ² /g with the same value.

From the top down, the first 175 cm is thermostated by means of a salt bath A countercurrently pumped with the reaction gas B. The second 175 cm is thermostated by means of a salt bath B which is countercurrent to the reaction gas B.

The second oxidation stage in the flow direction of the reaction gas B is also a multi-tube reactor having two temperature zones. The reaction tubes were loaded from the top down as follows: 1 section: length 20 cm as a block talc ring with an upstream bed geometry of 7 mm x 7 mm x 4 mm (outer diameter x length x inner diameter).

2 sections: length 90 cm 25 wt% (or 30 wt%) geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) block talc ring and 75% by weight (or 70% by weight) A catalyst filler from a homogeneous mixture of 4 sections of coated catalyst.

3 sections: length 50 cm 15% by weight (or 20% by weight) geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) block talc ring and 85% by weight (or 80% by weight) A catalyst filler from a homogeneous mixture of 4 sections of coated catalyst.

4 section: length 190 cm according to DE-A 100 46 928 Preparation Example 5 Ring (7 mm × 3 mm × 4 mm = outer diameter × length × inner diameter) coated catalyst catalyst filler (stoichiometry: Mo ₁₂ V ₃ W _1.2 Cu _2.4 O _x ).

From the top down, the first 175 cm is thermostated by means of a salt bath C that is countercurrent to the reaction gas. The second 175 cm is thermostated by means of a salt bath D which is countercurrently drawn with the reaction gas.

In both oxidation stages, the salt baths are each guided around the reaction tube in a curved manner due to the deflection plates.

A tube bundle heat exchanger cooled by means of a salt bath is placed between the two oxidation stages, and the product gas of the first oxidation stage can be cooled therewith. A valve for supplying molecular oxygen (purity > 99% by volume) is placed upstream of the second oxidation stage.

The propylene loading of the catalyst packing in the first oxidation stage is selected to be 130 l(STP)/l. h. The salt melt (53% by weight of KNO ₃ , 40% by weight of NaNO ₂ , 7% by weight of NaNO ₃ ) has the following inlet temperatures: T _A = 324 ° C T _B = 328 ° C T _{C =} 265 ° C T _D = 269 ° C .

Sufficient molecular oxygen (139 ° C, 2.33 bar) was metered (throttled) into the product gas mixture of the first oxidation stage, and the O ₂ : acrolein molar ratio in the resulting gas mixture of the second oxidation stage was 0.73. The acrolein loading of the catalyst packing in the second oxidation stage is 110 l(STP)/l. h. The pressure at the inlet to the second oxidation stage was 1.57 bar. The reaction gas B leaves the intercooler at a temperature of 260 ° C and the inlet temperature of the charged gas mixture into the second oxidation stage is 256 ° C. The product gas mixture of the first oxidation stage has the following contents: adialdehyde 16.84% by volume, acrylic acid 1.13 vol%, propane 36.88 vol%, propylene 0.99 vol%, methane 2.99 vol%, H ₂ 4.61 vol%, ethylene 0.99 vol% O ₂ 4.32 vol%, H ₂ O 23.96 vol%, CO 0.847 vol%, and CO ₂ 4.93 vol%.

The temperature of the product gas in the first oxidation stage was 335 ° C before entering the aftercooler.

The product gas mixture (product gas B) of the second oxidation stage has a temperature of 270 ° C and a pressure of 1.55 bar, and the following contents: 0.089 vol% of acrolein, 17.02 vol% of acrylic acid, 0.46 vol% of acetic acid, and 36.79 vol% of propane MeOH 0.99 vol%, methane 2.98 vol%, ethylene 0.99 vol%, H ₂ 4.59 vol%, O ₂ 2.69 vol%, H ₂ O 24.57 vol%, CO 1.35 vol% and CO ₂ 5.91 vol%.

As described in WO 2004/035514, product gas B (5.15 m ³ (STP) / h) is fractionally condensed in a tray column (separation zone II).

As the first fuel, 13.7 g/h of a high boiling point (polyacrylic acid (Michael adduct) or the like) was sent to the residue for incineration.

From the second collection tray above the feed of product gas B in the tray column, 2.80 kg/h of condensed crude acrylic acid having a temperature of 15 ° C and 96.99 wt% of acrylic acid was withdrawn. As described in WO 2004/035514, it is suspended and crystallized after the addition of a small amount of water, the suspended crystals are separated from the mother liquor in a hydraulic scrubber and the mother liquor is recycled to the condensation column as described in WO 2004/035514. The washed suspension crystals have a purity of >99.87 wt% acrylic acid and are suitable for the immediate preparation of polymers for "superabsorbent" water in sanitary articles.

The amount of acid water condensate drawn from the third collection tray above the feed to the product gas B of the condensation column but not recycled to the condensation column is 1.10 kg/h, has a temperature of 33 ° C and has the following contents : 0.80% by weight of acrolein, 4.92% by weight of acrylic acid, 5.73% by weight of acetic acid and 87.0% by weight of water.

It is also sent to the residue for incineration (the air of 5.49 m ³ (STP) / h is sent as an oxygen source to the residue for incineration).

At the top of the condensation column, the residual gas of 2.98 m ³ (STP) / h leaves the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar and below: acryl aldehyde 0.03 vol%, acrylic acid 0.02 vol%, acetic acid 0.01 % by volume, propane 63.7 vol%, propylene 1.72 vol%, H ₂ 6.90 vol%, methane 5.16 vol%, ethylene 1.72 vol%, O ₂ 4.68 vol%, H ₂ O 2.07 vol%, CO 2.34 vol% and CO ₂ 10.18 volume%.

In the first compressor stage of the multistage radial compressor, the residual gas is compressed from 1.20 bar to 2.40 bar, during which the temperature of the residual gas rises to 70 °C. Subsequently, the compressed residual gas is divided into two portions having the same composition. The first part (1.19 m ³ (STP)/h) is formed as a filling gas mixture which is supplied to the reaction zone A as a (compressed) residual gas recycle gas. The other part (1.79 m ³ (STP) / h) is compressed from 2.40 bar to 6.93 bar in the second compressor stage. This heats the portion to 130 °C.

In the indirect heat exchanger, the residual gas is cooled to 113 ° C without condensate formation (the coolant is "propane washed" residual gas (which is also referred to as "exhaust gas" hereinafter), in the indirect heat exchanger from 30 After heating to 86 ° C, it decompresses from 20 bar in the first expansion stage of the multistage expansion turbine and simultaneously cools). This heated exhaust gas (0.58 m ³ (STP) / h). In the second expansion stage, the exhaust gas is depressurized to 1.10 bar (which subsequently has a temperature of 25 ° C) and then sent to the residue for incineration.

The compressed residual gas (P = 6.93 bar and T = 113 ° C) was cooled to 59 ° C in an indirect air cooler. In the third compressor stage, it is compressed from 6.93 bar to 20 bar while heating it to 116 °C.

In the downstream indirect heat exchanger, the residual gas is cooled to 107 ° C (the coolant is "propane washed" residual gas, which is simultaneously heated from 30 ° C to 86 ° C) without condensate formation. In a subsequent further indirect heat exchanger, the residual gas at 107 ° C is cooled to 35 ° C (the coolant is surface water of about 25 ° C; it is also often possible to use other cooling water instead of surface water in this paper); it will be 29 g / The water of h is condensed from the residual gas. The aqueous condensate, which is preferably separated by a small droplet separator, is likewise sent to the residue for incineration and is incinerated therein along with the other residues mentioned (while, as already stated, supply 5.49 m ³ (STP)/ h air). The heat incineration gas was cooled to produce steam (52 bar, 267 ° C, 4.6 kg/h) and released into the environment.

The residual gas, which still contains 1.75 m ³ (STP) / h of propane (20 bar, 35 ° C), is sent to the lower section of the column with the random packing (separation zone III). At the top of the scrubber, there will be an introduction temperature of 30 ° C (P = 20 bar), and 12.96 kg / h of industrial tetradecane from PKWF 4/7 af of Haltermann, Germany as an absorbent (by means of The gas chromatographic analysis of F/D detection produced the following GC area % composition at the beginning (fresh): n-tridecane 7.6%, n-tetradecane 47.3%, n-pentadecane 7.0%, n-hexadecane 3.2 %, n-heptadecane 14.1% and total residue 20.7%.

At the top of the scrubber, the "propane-washed" residual gas (exhaust gas) escapes at a pressure of 20 bar and a temperature of 30 ° C and an amount of 0.58 m ³ (STP) / h, which has the following contents: Propane 1.97 vol%, propylene 0.049 vol%, methane 15.95 vol%, ethylene 5.32 vol%, H ₂ 21.29 vol%, O ₂ 14.46 vol%, H ₂ O 0.82 vol%, CO 7.23% by volume and CO ₂ 31.43 vol%.

The exhaust gas is subjected to indirect heat exchange as described and then sent to the residue for incineration or by means of flame incineration.

15.27 kg/h of absorbent having a pressure of 20 bar and a temperature of 59 ° C was withdrawn from the bottom of a "propane scrubber" which was not externally heated or cooled. It comprises 14.79% by weight of propane and 0.38% by weight of propylene.

It is heated to 69 ° C in an indirect heat exchanger before the absorber is directed to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (12.96 kg/h) which had a temperature of 80 ° C at a pressure of 3.26 bar and still included 0.09 wt% of dissolved propane. It was compressed back to 20 bar by means of a pump and after being indirectly cooled to 30 ° C with surface water (25 ° C), it was recycled to the top of the "propane scrubber" to propane absorption.

After heating to 69 ° C, the absorbent (for example in a reverse pump or by means of a valve) is depressurized to a pressure of 3.26 bar (the mechanical energy released under reverse pump conditions is also used to recompress the desorption tower) A propane-containing absorbent (liquid effluent from the desorption column)), and the obtained two-phase mixture is introduced into the desorption column at the top.

At a pressure of 5 bar and a temperature of 152 ° C, 0.1 kg / h of steam as a stripping gas from the bottom and the absorption descending from the top of the desorption tower is sent back to the desorption tower (the same tower with random packing) ).

Disposed in the desorption column is a heating coil, which is also guided to have a temperature of 152 ° C, P = 5.00 bar, 0.5 kg / h of steam through the heating coil and the rising vapor in the desorption tower countercurrent.

At the top of the desorption column, 1.28 m ³ (STP)/h of propane recycle gas having a temperature of 69 ° C and a working pressure of 3.26 bar was released. After heating to 400 ° C by heat exchange with the product gas A having a temperature of T = 435 ° C, the propane cycle gas is used as an actuating injection to operate the jet pump, and the dehydrogenation cycle gas and the propane recycle gas system are pumped by the jet Filling to reaction zone A using the starting mixture of reaction gas A.

The propane recycle gas had the following contents: propane 88.45 vol%, propylene 2.39 vol%, and H ₂ O 8.66% by volume.

III. Second illustrative method for preparing acrylic acid from propane (describes stable operating conditions)

Reaction zone A is a tray reactor as described in II. The heterogeneously catalyzed partial propane dehydrogenation is carried out in the tray reactor in a direct transfer mode with a dehydrogenation recycle gas. The total amount of catalyst in all disks versus propane loading (calculated without inert material) is 1500 l(STP)/l. h.

8.23 m ³ (STP) / h of the reaction gas A filling gas mixture (starting mixture: P = 1.84 bar) is sent to the first catalyst bed in the flow direction and has the following contents: propane 64.20% by volume, propylene 16.0 volume %, methane 1.39 vol%, ethylene 0.46 vol%, H ₂ 2.76 vol%, O ₂ 1.06 vol% and H ₂ O 13.13 vol%.

The reaction gas A is filled with a gas mixture as a mixture of a first (1.15 m ³ (STP) / h) and a second (7.08 m ³ (STP) / h) gas mixture by combining two gases by means of a static mixer Obtained from the mixture. The first gas mixture has a temperature of 450 ° C and a pressure of 1.84 bar and comprises: 1.06 m ³ (STP) / h of crude propane comprising propane (fresh propane) to a level of > 98 vol% (T = 237 ° C, P = 2.90 bar), and 0.07 m ³ (STP) / h of O ₂ (T = 143 ° C, P = 1.84 bar, purity > 99 vol%).

The indirect heat exchange with the product gas A in the first heat exchanger with the derivation reaction zone A and in the direction of the reaction zone B produces a temperature of 450 °C. The second gas mixture comprises: 5.04 m ³ (STP) / h of dehydrogenation cycle gas (T = 540 ° C, P = 1.73 bar) and 2.04 m ³ (STP) / h of propane cycle gas (T = 400 ° C, P = 3.26 bar).

It has a pressure of 1.84 bar and a temperature of 486 °C. The indirect heat exchange with the product gas A cooled to a temperature of 476 ° C in the first heat exchanger produces a temperature of the propane recycle gas. The dehydrogenation cycle gas stream is achieved according to the principle of a jet pump (cf. DE-A 10211275), and the propane cycle gas heated to 400 ° C acts as an actuating jet.

The inlet temperature of the reaction gas A charging gas mixture and the bed height of the first catalyst bed through which the gas mixture is filled by the reaction gas A are adjusted so that the reaction gas A having the following contents leaves at a temperature of 503 ° C and a pressure of 1.81 bar. Catalyst bed: propane 60.7 vol%, propylene 18.19 vol%, methane 1.59 vol%, ethylene 0.53 vol%, H ₂ 2.94 vol%, O ₂ vol%, and H ₂ O 15.0 vol%.

The amount of departure is 8.35 m ³ (STP) / h.

When the first catalyst bed was passed, 0.13 m ³ (STP) / h of molecular oxygen (purity > 99 vol%) was metered into the reaction gas A (T = 143 ° C, P = 2.40 bar). The metered addition was achieved by throttling so that the resulting pressure of the resulting reaction gas A was still 1.81 bar.

The bed height of the second catalyst bed was such that the reaction gas A having the following contents left the second catalyst bed at a temperature of 520 ° C and a pressure of 1.78 bar: propane 54.2% by volume, propylene 21.3% by volume, methane 1.90% by volume, ethylene. 0.63 vol%, H ₂ 3.61 vol%, O ₂ vol%, and H ₂ O 17.3 vol%.

The amount of departure is 8.70 m ³ (STP) / h.

Upstream of the static mixer placed upstream of the third catalyst bed, 0.17 m ³ (STP) / h of molecular oxygen (purity > 99 vol%) was metered into the reaction gas A (T = 143 ° C, at 1.78 Throttling under the pressure of Ba). The resulting pressure of the resulting reaction gas was still 1.78 bar.

The bed height of the third catalyst bed was such that the reaction gas A exited the third catalyst bed as a product gas A having the following contents at a temperature of 540 ° C and a pressure of 1.73 bar: propane 46.30 vol%, propylene 25.12 vol%, ethylene 0.75. % by volume, methane 2.27 vol%, H ₂ 4.50 vol%, O ₂ 0 vol%, and H ₂ O ₂ 0.05 vol%.

The amount of departure is 9.17 m ³ (STP) / h.

The splitter divides the product gas A into two parts having the same composition. The first part is 5.04 m ³ (STP)/h and the second part is 4.13 m ³ (STP)/h.

The first portion is recycled to the reaction zone A as a component of the reaction gas A charging gas mixture. For the purpose of the first part of the dehydrogenation recycle gas stream, the heated propane recycle gas is used as an actuating injection to operate the jet pump, and the transport direction of the decelerated actuated jet passes through the mixing zone through the actuating nozzle, And the diffuser is directed in the direction of the inlet of the reaction zone A, and the nozzle is directed in the direction of the first flow portion of the product gas A, and the "nozzle-mixing zone-diffuser" is connected to form the first part of the product gas A to be recycled The only connecting line between the picking up of reaction zone A.

The second portion of product gas A is directed to reaction zone A and first cooled from 540 ° C to 476 ° C by a first indirect heat exchange with the first gas mixture, the first gas mixture comprising crude propane and molecular oxygen, included in The reaction gas A is initially mixed (filled with a gas mixture). The working pressure of the product gas A is substantially maintained.

At the same time, the first gas mixture is heated from 235 ° C to 450 ° C while substantially maintaining a working pressure of 1.84 bar of the gas mixture.

Product gas A was cooled from 476 ° C to 307 ° C in a second indirect heat exchange downstream of the propane recycle gas from separation zone IV. Conversely, the temperature of the propane recycle gas increased from 69 ° C to 400 ° C.

Subsequently, 0.52 m ³ (STP) / h of steam was removed by condensation of product gas A from cooling to 307 ° C. To this end, the product gas A, which is cooled to 307 ° C, is indirectly indirectly with a suitable amount of steam which has been removed by condensation to a product gas A ^* (T = 54 ° C, P = 1.73 bar) of 3.60 m ³ (STP) / h. Heat exchange, and therefore cooled to 111 °C. The temperature of the product gas A ^* was simultaneously increased to 277 °C.

Direct cooling of the atomized aqueous condensate that has been cooled (to 29 ° C) and pre-removed from product gas A by suitable direct cooling removes the above amount of steam from product gas A by condensation to obtain product gas A. ^* (P = 1.73 bar, T = 54 ° C, 3.60 m ³ (STP) / h). It was heated to 277 ° C as described above. Molecular oxygen (purity > 99% by volume, T = 143 ° C, P = 2.40 bar) of 1.39 m ³ (STP) / h is then throttling into the product gas A ^* . It forms a reaction gas B starting mixture (filling gas mixture) having a temperature of 262 ° C and an operating pressure of 1.63 bar.

The reaction gas B starting mixture (filling gas mixture) has the following contents: propane 38.26 vol%, propylene 20.77 vol%, ethylene 0.62 vol%, methane 1.87 vol%, H ₂ 3.77% by volume, O ₂ 27.53 vol%, and H ₂ O 6.20% by volume.

It is used to load the reaction zone B having the same structure as that of the structure from II.

The propylene loading of the catalyst packing in the first oxidation stage was selected to be 164 l(STP)/l. h. The salt melt (53% by weight of KNO ₃ , 40% by weight of NaNO ₂ , 7% by weight of NaNO ₃ ) has the following inlet temperatures: T _A = 331 ° C T _B = 339 ° C T _C = 275 ° C T _D = 282 ° C .

Sufficient molecular oxygen (143 ° C, 2.40 bar) was metered (throttled) into the product gas mixture of the first oxidation stage, and the O ₂ : acrolein molar ratio in the resulting gas mixture of the second oxidation stage was 0.72. The acrolein loading of the catalyst packing in the second oxidation stage is 139 l(STP)/l. h. The pressure at the inlet to the second oxidation stage was 1.57 bar.

The reaction gas B leaves the intercooler at a temperature of 260 ° C and the inlet temperature of the charged gas mixture into the second oxidation stage is 256 ° C.

The product gas mixture of the first oxidation stage has the following contents: acryl amide 17.51 vol%, acrylic acid 1.18 vol%, propane 38.12 vol%, propylene 1.03 vol%, methane 1.86 vol%, H ₂ 3.71 vol%, ethylene 0.62 vol% O ₂ 4.18 vol%, H ₂ O 27.83 vol%, CO 0.88 vol%, and CO ₂ 2.07 vol%.

The temperature of the product gas in the first oxidation stage was 335 ° C before entering the aftercooler.

The product gas mixture (product gas B) of the second oxidation stage has a temperature of 270 ° C and a pressure of 1.55 bar, and the following contents: 0.089 vol% of acrolein, 17.67 vol% of acrylic acid, 0.52 vol% of acetic acid, and 37.99 vol% of propane 1.02 vol% of propylene, 1.86 vol% of methane, 0.62 vol% of ethylene, 3.70 vol% of H ₂ , 2.61 vol% of O ₂ , 28.42 vol% of H ₂ O, 1.41 vol% of CO and 3.10 vol% of CO ₂ .

As described in WO 2004/035514, product gas B (5.03 m ³ (STP) / h) is fractionally condensed in a tray column (separation zone II).

As the first fuel, a high boiling point of 13.9 g/h (polyacrylic acid (Michael adduct) or the like) was sent to the residue for incineration.

From the second collection tray above the feed of product gas B in the tray column, 2.82 kg/h of condensed crude acrylic acid having a temperature of 15 ° C and 96.99 wt% of acrylic acid was withdrawn. As described in WO 2004/035514, it is suspended and crystallized after the addition of a small amount of water, the suspended crystals are separated from the mother liquor in a hydraulic scrubber and the mother liquor is recycled to the condensation column as described in WO 2004/035514. The washed suspension crystals have a purity of >99.87 wt% acrylic acid and are suitable for the immediate preparation of polymers for "superabsorbent" water in sanitary articles.

The amount of acid water condensate drawn from the third collection tray above the feed to the product gas B of the condensation column but not recycled to the condensation column is 1.24 kg/h, has a temperature of 33 ° C and has the following inclusions : 0.73 wt% acrolein, 4.95 wt% acrylic acid, 5.16 wt% acetic acid, and 88.1 wt% water.

At the top of the condensation column, the residual gas of 2.72 m ³ (STP) / h leaves the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar and below: acryl aldehyde 0.03 vol%, acrylic acid 0.02 vol%, acetic acid 0.01 % by volume, propane 70.26 vol%, propylene 1.91 vol%, ethylene 1.15 vol%, H ₂ 6.93% by volume, methane 3.44 vol%, O ₂ 4.83 vol%, H ₂ O 2.11 vol%, CO 2.59 vol% and CO ₂ 5.73 volume%.

In the first compressor stage of the multistage radial compressor, the residual gas is compressed from 1.20 bar to 6.93 bar, during which the temperature of the residual gas rises to 127 °C.

In the indirect heat exchanger, the residual gas is cooled to 114 ° C without condensate formation (the coolant is "propane washed" residual gas (which is also referred to hereinafter as "exhaust gas"), in the indirect heat exchanger from 30 After heating to ° °C, °C begins to decompress from 20 bar in the first expansion stage of the multistage expansion turbine and simultaneously cools). This heated exhaust gas (0.69 m ³ (STP) / h). In the second expansion stage, the exhaust gas is depressurized to 1.10 bar (which subsequently has a temperature of 18 ° C) and then sent to the residue for incineration.

The compressed residual gas (P = 6.93 bar and T = 114 ° C) was cooled to 59 ° C in an indirect air cooler. In another compressor stage, it is compressed from 6.93 bar to 20 bar while heating it to 114 °C. In the downstream indirect heat exchanger, it is cooled to 107 ° C (the coolant is "propane washed" residual gas, which is simultaneously heated from 30 ° C to 84 ° C) without condensate formation. In the downstream of another indirect heat exchanger, the residual gas at 107 ° C was cooled to 35 ° C (the coolant was surface water of about 25 ° C).

It condenses 44.8 g/h of water from the residual gas. The aqueous condensate is likewise sent to the residue for incineration and is incinerated therein along with the other residues mentioned (while supplying 5.95 m ³ (STP) / h of air). The hot combustion gases were cooled to obtain steam (52 bar, 267 ° C, 5.16 kg/h) and released into the environment.

The residual gas, which still contains 2.67 m ³ (STP) / h of propane (20 bar, 35 ° C), was introduced into the lower section of the column (separation zone III) with the random packing.

At the top of the scrubber, 20.39 kg/h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 °C (P = 20 bar).

At the top of the scrubber, the "propane-washed" residual gas (exhaust gas) escapes at a pressure of 20 bar and a temperature of 30 ° C and an amount of 0.69 m ³ (STP) / h, which has the following contents: Propane 1.38 vol%, propylene 0.04 vol%, ethylene 4.48% by volume, methane 13.44 vol%, H ₂ 27.12 vol%, O ₂ 18.89 vol%, H ₂ O 1.04 vol%, CO 10.16 vol% and CO ₂ 22.44 vol%.

The exhaust gas is subjected to indirect heat exchange as described and then sent to the residue for incineration or by means of flame incineration.

24.27 kg/h of absorbent having a pressure of 20 bar and a temperature of 61 ° C was withdrawn from the bottom of a "propane scrubber" which was not externally cooled or heated. It comprises 15.63% by weight of propane and 0.41% by weight of propylene.

It is heated to 69 ° C in an indirect heat exchanger before the absorber is directed to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (20.39 kg/h) which had a temperature of 80 ° C at a pressure of 3.26 bar and still included 0.089 wt% of dissolved propane. It was compressed back to 20 bar by means of a pump and after being indirectly cooled to 30 ° C with surface water (25 ° C), it was recycled to the top of the "propane scrubber" to propane absorption.

After heating to 69 ° C, the absorbent (for example in a reverse pump or by means of a valve) is depressurized to a pressure of 3.26 bar (the mechanical energy released under reverse pump conditions is also used to recompress the desorption tower) A propane-containing absorbent (liquid effluent from the desorption column)), and the obtained two-phase mixture is introduced into the desorption column at the top.

At a pressure of 5 bar and a temperature of 152 ° C, 0.1 kg / h of steam as a stripping gas from the bottom and the absorption descending from the top of the desorption tower is sent back to the desorption tower (the same tower with random packing) ).

Disposed in the desorption column is a heating coil, which is also guided to have a temperature of 152 ° C, P = 5.00 bar, 0.8 kg / h of steam through the heating coil and the rising vapor in the desorption tower countercurrent.

At the top of the desorption column, a propane recycle gas of 2.04 m ³ (STP)/h having a temperature of 69 ° C and a working pressure of 3.26 bar escaped. After heating to 400 ° C by heat exchange with the product gas A having a temperature of T = 476 ° C, the propane recycle gas is used as an actuating spray to operate the jet pump, and the dehydrogenation cycle gas and the propane recycle gas system in the mixture are used. The jet pump is pumped to the reaction zone A using the starting mixture of the reaction gas A.

The propane cycle gas had the following contents: propane 93.07 vol%, propylene 2.52 vol%, and H ₂ O 3.40 vol%.

IV. Third illustrative method for preparing acrylic acid from propane (describes stable operating conditions)

Reaction zone A is a tray reactor as described in II.

The heterogeneously catalyzed partial propane dehydrogenation is carried out in the disc reactor in a direct transfer mode without dehydrogenation recycle gas. The total amount of catalyst in all disks versus propane loading (calculated without inert material) is 1500 l(STP)/l. h.

3.80 m ³ (STP) / h of the reaction gas A filling gas mixture (starting mixture: T = 418 ° C; P = 1.77 bar) is sent to the first catalyst bed in the flow direction and has the following content: propane 75.21 volume %, propylene 1.35 vol%, methane 0.01 vol%, H ₂ 8.60 vol%, O ₂ 4.32 vol%, N ₂ 0.96 vol%, H ₂ O 7.68 vol%, CO 0.02 vol%, and CO ₂ 0.83 vol%.

The reaction gas A is filled with a gas mixture as a mixture of a first (1.59 m ³ (STP) / h) and a second (2.21 m ³ (STP) / h) gas mixture by combining two gases by means of a static mixer Obtained from the mixture. The first gas mixture has a temperature of 450 ° C and a pressure of 1.77 bar and comprises: 1.02 m ³ (STP) / h of crude propane comprising propane (fresh propane) to a level of > 98 vol% (T = 237 ° C, P = 2.90 bar), 0.16 m ³ (STP) / h of O ₂ (T = 143 ° C, P = 1.77 bar, purity > 99 vol%), and 0.41 m ³ (STP) / h of gas, which mainly includes Molecular hydrogen was removed from the "propane-washed" residual gas (exhaust gas) by membrane separation (T = 30 ° C, which throttled to 1.77 bar).

The indirect heat exchange with the product gas A in the first heat exchanger with the derivation reaction zone A and in the direction of the reaction zone B produces a temperature of 450 °C.

The second gas mixture was a propane recycle gas (2.21 m ³ (STP) / h; 400 ° C; 1.77 bar). The indirect heat exchange with the product gas A cooled to a temperature of 500 ° C in the first heat exchanger produces a temperature of the propane recycle gas.

Adjusting the inlet temperature of the reaction gas A charging gas mixture and the bed height of the first catalyst bed through which the gas mixture is filled by the reaction gas A so that the reaction gas A having the following contents leaves at a temperature of 507 ° C and a pressure of 1.76 bar Catalyst bed: propane 64.42 vol%, propylene 9.10 vol%, ethylene 0.079 vol%, methane 0.33 vol%, H ₂ 7.66% by volume, O ₂ 0 vol%, N ₂ 0.92 vol%, H ₂ O 15.69 vol%, CO 0 % by volume, and CO ₂ 0.82% by volume.

The amount of departure is 3.95 m ³ (STP) / h.

When the first catalyst bed was passed, 0.13 m ³ (STP) / h of molecular oxygen (purity > 99 vol%) was metered into the reaction gas A (T = 143 ° C, P = 2.40 bar). The metered addition was achieved by throttling so that the resulting pressure of the resulting reaction gas A was still 1.76 bar.

The bed height of the second catalyst bed was such that the reaction gas A having the following contents left the second catalyst bed at a temperature of 544 ° C and a pressure of 1.75 bar: propane 51.72 vol%, propylene 15.97 vol%, ethylene 0.149 vol%, methane 0.60 vol%, H ₂ 8.68 vol%, O ₂ 0 vol%, N ₂ 0.85 vol%, H ₂ O 20.26 vol%, CO 0 vol%, and CO ₂ 0.76 vol%.

The amount of departure is 4.29 m ³ (STP) / h.

Upstream of the static mixer placed upstream of the third catalyst bed, 0.12 m ³ (STP) / h of molecular oxygen (purity > 99 vol%) was metered into the reaction gas A (T = 143 ° C, at 1.75 Throttling under the pressure of Ba). The resulting pressure of the resulting reaction gas was still 1.75 bar.

The bed height of the third catalyst bed was such that the reaction gas A exited the third catalyst bed as a product gas A having the following contents at a temperature of 571 ° C and a pressure of 1.73 bar: propane 40.12 vol%, propylene 22.2 vol%, ethylene 0.22. % by volume, methane 0.87 vol%, H ₂ 10.11 vol%, O ₂ 0 vol%, N ₂ 0.79 vol%, H ₂ O 23.96 vol%, CO 0 vol%, and CO ₂ 0.70 vol%.

The product gas A is led to the reaction zone A and first cooled from 571 ° C to 500 ° C by a first indirect heat exchange with the first gas mixture, the first gas mixture being contained in the starting mixture of the reaction gas A (filling gas mixture) It also includes crude propane, a gas including molecular hydrogen, and molecular oxygen. The working pressure of the product gas A is substantially maintained. This heats the first gas mixture from 213 ° C to 450 ° C while substantially maintaining the working pressure of the gas mixture of 1.77 bar.

The product gas A was cooled from 500 ° C to 337 ° C in a subsequent second indirect heat exchange with the propane recycle gas from separation zone IV. Conversely, the temperature of the propane recycle gas increased from 69 ° C to 400 ° C.

Subsequently, 0.79 m ³ (STP) / h of steam was separated by condensation of product gas A from cooling to 337 ° C. To this end, the product gas A, which is cooled to 307 ° C, is first with a suitable amount of steam which has been previously removed by condensation to a product gas A ^* of 3.85 m ³ (STP) / h (T = 54 ° C; P = 1.73 bar) Indirect heat exchange, and therefore cooled to 120 °C.

The temperature of the product gas A ^* was simultaneously increased to 307 °C.

Direct cooling of the atomized aqueous condensate that has been cooled (to 29 ° C) and pre-removed from product gas A by suitable direct cooling removes the above amount of steam from product gas A by condensation to obtain product gas A. ^* (P = 1.73 bar, T = 54 ° C, 3.85 m ³ (STP) / h).

It was heated to 307 ° C as described above. Molecular oxygen (purity > 99% by volume, T = 143 ° C, P = 2.40 bar) of 1.41 m ³ (STP) / h was then throttling into the product gas A ^* . It forms a reaction gas B starting mixture (filling gas mixture) having a temperature of 290 ° C and an operating pressure of 1.63 bar.

The starting mixture of reaction gas B (filling gas mixture) had the following contents: propane 35.35 vol%, propylene 19.60 vol%, ethylene 0.19 vol%, methane 0.76 vol%, H ₂ 8.91 vol%, O ₂ 26.60 vol% N ₂ 0.69 vol%, H ₂ O 6.27 vol%, CO 0 vol%, and CO ₂ 0.62 vol%.

It is used to load the reaction zone B having the same structure as that of the structure from II.

The propylene loading of the catalyst packing in the first oxidation stage was selected to be 1301 (STP) / l. h. The salt melt (53% by weight of KNO ₃ , 40% by weight of NaNO ₂ , 7% by weight of NaNO ₃ ) has the following inlet temperatures: T _A = 322 ° C T _B = 327 ° C T _C = 261 ° C T _D = 268 ° C .

Sufficient molecular oxygen (143 ° C, 2.40 bar) was metered (throttled) into the product gas mixture of the first oxidation stage, and the O ₂ : acrolein molar ratio in the resulting gas mixture of the second oxidation stage was 0.745. The acrolein loading of the catalyst packing in the second oxidation stage is 110 1 (STP) / l. h. The pressure at the inlet to the second oxidation stage was 1.57 bar. The reaction gas B leaves the intercooler at a temperature of 260 ° C and the inlet temperature of the charged gas mixture into the second oxidation stage is 256 ° C.

The product gas mixture of the first oxidation stage has the following contents: acrolein 16.51 vol%, acrylic acid 1.11 vol%, propane 35.23 vol%, propylene 0.98 vol%, ethylene 0.19 vol%, methane 0.76 vol%, H ₂ 8.8 vol% O ₂ 4.57% by volume, N ₂ 0.69 vol%, H ₂ O 26.66 vol%, CO 0.83 vol%, and CO ₂ 2.56 vol%.

The temperature of the product gas in the first oxidation stage was 335 ° C before entering the aftercooler.

The product gas mixture of the second oxidation stage (product gas B) has a temperature of 270 ° C and a pressure of 1.55 bar, and the following contents: acrolein 0.08 vol%, acrylic acid 16.71 vol%, acetic acid 0.495 vol%, propane 35.20 vol% , propylene 0.89 vol%, ethylene 0.19 vol%, methane 0.76 vol%, H ₂ 8.87 vol%, O ₂ 2.86 vol%, N ₂ 0.69 vol%, H ₂ O 27.28 vol%, CO 1.33 vol%, and CO ₂ 3.54 volume%.

Product gas B (5.29 m ³ (STP) / h) was fractionally condensed in a tray column (separation zone II) as described in WO 2004/035514.

As the first fuel, 13.8 g/h of a high boiling point (polyacrylic acid (Michael adduct) or the like) was sent to the residue for incineration.

From the second collection tray above the feed of product gas B in the tray column, 2.80 kg/h of condensed crude acrylic acid having a temperature of 15 ° C and 96.99 wt% of acrylic acid was withdrawn. As described in WO 2004/035514, it is suspended and crystallized after the addition of a small amount of water, the suspended crystals are separated from the mother liquor in a hydraulic scrubber and the mother liquor is recycled to the condensation column as described in WO 2004/035514.

The washed suspension crystals have a purity of >99.87 wt% acrylic acid and are suitable for the immediate preparation of polymers for "superabsorbent" water in sanitary articles.

The amount of acid water condensate drawn from the third collection tray above the feed to the product gas B of the condensation column but not recycled to the condensation column is 1.25 kg/h, has a temperature of 33 ° C and has the following inclusions : 0.72% by weight of acrolein, 4.95 % by weight of acrylic acid, 5.09% by weight of acetic acid and 88.24% by weight of water.

At the top of the condensation column, 2.98 m ³ (STP) / h of residual gas leaves the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar and below: acryl 0.02 vol%, acrylic acid 0.02 vol%, propane 62.39 % by volume, 1.73% by volume of propylene, 0.34% by volume of ethylene, 1.35 vol% of methane, H ₂ 16.23 vol%, O ₂ 5.05 vol%, N ₂ 1.27 vol%, H ₂ O 1.92 vol%, CO 2.36 vol%, and CO ₂ 6.33 vol%.

In the first compressor stage of the multistage radial compressor, the residual gas is compressed from 1.20 bar to 6.93 bar, during which the temperature of the residual gas rises to 134 °C.

In the indirect heat exchanger, the residual gas is cooled to 123 ° C without condensate formation (the coolant is "propane washed" residual gas (which is also referred to as "exhaust gas" hereinafter), and the molecules are removed by means of a separation membrane Hydrogen and after heating from 30 ° C to 102 ° C in an indirect heat exchanger, it decompresses from 20 bar in the first expansion stage of the multistage expansion turbine and simultaneously cools). This heated exhaust gas (0.59 m ³ (STP) / h). In the second expansion stage, the exhaust gas is depressurized to 1.10 bar (which subsequently has a temperature of 23 ° C) and then sent to the residue for incineration.

The compressed residual gas (P = 6.93 bar and T = 134 ° C) was cooled to 59 ° C in an indirect air cooler. In another compressor stage, it is compressed from 6.93 bar to 20 bar while heating it to 117 °C. In the downstream indirect heat exchanger, it is cooled to 110 ° C (the coolant is "propane washed" residual gas, when the H ₂ film removal is completed, it is simultaneously heated from 30 ° C to 102 ° C) without condensate formation . In the downstream of another indirect heat exchanger, the residual gas at 110 ° C is cooled to 35 ° C (the coolant is surface water of about 25 ° C).

It condenses 44.2 g/h of water from the residual gas. The aqueous condensate is likewise sent to the residue for incineration and is incinerated together with the other residues mentioned while supplying 4.12 m ³ STP) / h of air. The hot combustion gases were cooled to obtain steam (52 bar, 267 ° C, 3.41 kg/h) and released into the environment.

The residual gas, which still contains 2.93 m ³ (STP) / h of propane (20 bar, 35 ° C), was introduced into the lower section of the column (separation zone III) with the random packing. At the top of the scrubber, 21.62 kg/h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 °C (P = 20 bar).

At the top of the scrubber, the "propane-washed" residual gas (exhaust gas) escapes at a pressure of 20 bar and a temperature of 30 ° C and an amount of 1.01 m ³ (STP) / h, which has the following inclusions: Propane 0.92% by volume, propylene 0.03% by volume, ethylene 0.99 vol%, methane 3.99 vol%, H ₂ 47.97 vol%, O ₂ 14.94 vol%, N ₂ 3.75 vol%, H ₂ O 0.78 vol%, CO 6.95% by volume and CO ₂ 18.69 vol%.

The "propane-washed" off-gas was directed to a bundle of cast film (external pressure: 1.77 bar; B-H polyimine film from UBE Industries Ltd.) and left (P = 20 bar; T = 30) °C), simultaneously releasing 0.41 m ³ (STP) / h of gas stream (permeate stream) with the following contents: ethylene 0.02 vol%, methane 0.059 vol%, H ₂ 79.15 vol%, O ₂ 1.53 vol%, N ₂ 8.83 vol%, H ₂ O 1.52 vol%, CO 0.22 vol%, and CO ₂ 7.68% by volume.

The gas stream is sent (throttled) to the first gas mixture to produce a reactive gas A charge gas mixture.

The remaining exhaust gas is decompressed as described and subjected to indirect heat exchange and then sent to the residue for incineration or incineration by flame.

From the bottom of the "propane scrubber" which was not externally cooled or heated, 25.40 kg/h of absorbent having a pressure of 20 bar and a temperature of 59 ° C was withdrawn. It comprises 14.40% by weight of propane and 0.38% by weight of propylene.

It is heated to 69 ° C in an indirect heat exchanger before the absorber is directed to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (21.62 kg/h) which had a temperature of 80 ° C at a pressure of 3.26 bar and still included 0.09 wt% of dissolved propane. It was compressed back to 20 bar by means of a pump and after being indirectly cooled to 30 ° C with surface water (25 ° C), it was recycled to the top of the "propane scrubber" to propane absorption.

After heating to 69 ° C, the absorbent (for example in a reverse pump or by means of a valve) is depressurized to a pressure of 3.26 bar (the mechanical energy released under reverse pump conditions is also used to recompress the desorption tower) A propane-containing absorbent (liquid effluent from the desorption column)), and the obtained two-phase mixture is introduced into the desorption column at the top.

At a pressure of 5 bar and a temperature of 152 ° C, 0.2 kg / h of steam as a stripping gas from the bottom and the absorption descending from the top of the desorption tower is sent back to the desorption tower (the same tower with random packing) ).

Disposed in the desorption column is a heating coil, which is also guided to have a temperature of 152 ° C, P = 5.00 bar, 0.8 kg / h of steam through the heating coil and the rising vapor in the desorption tower countercurrent.

At the top of the desorption column, 2.21 m ³ (STP)/h of propane recycle gas having a temperature of 69 ° C and a working pressure of 3.26 bar escaped.

After heating to 400 ° C by heat exchange with product gas A at a temperature T = 476 ° C, the propane recycle gas is sent to the reaction zone A using the starting mixture of the reaction gas A.

The propane recycle gas had the following contents: propane 83.73 vol%, propylene 2.32 vol%, and H ₂ O 12.94 vol%.

V. Fourth illustrative method for preparing acrylic acid from propane (describes stable operating conditions)

Reaction zone A is a tray reactor as described in II, with the exception that in the present example, downstream of it is another adiabatic blast furnace reactor having only one fixed catalyst bed and in its configuration Corresponding to the first fixed bed in the disc reactor in the flow direction. An indirect heat exchanger is connected between the two reactors. The heterogeneously catalyzed partial propane dehydrogenation is carried out in a tray reactor in a direct transfer mode with a dehydrogenation recycle gas. The total amount of catalyst in all disks versus propane loading (calculated without inert material) is 1500 l(STP)/l. h. In the downstream blast furnace reactor, substantially only the heterogeneous phase catalytic combustion of molecular hydrogen is achieved.

6.93 m ³ (STP) / h of the reaction gas A filling gas mixture (starting mixture: P = 2.46 bar) is sent to the first catalyst bed of the disk reactor in the flow direction and has the following content: propane 60.54 volume %, propylene 10.38 vol%, ethylene 0.28 vol%, methane 0.87 vol%, H ₂ 2.44 vol%, O ₂ 1.19 vol%, N ₂ 16.48 vol%, H ₂ O 6.82 vol%, CO 0 vol%, and CO ₂ 0.01% by volume.

The reaction gas A is filled with a gas mixture as a mixture of a first (1.47 m ³ (STP) / h) and a second (5.46 m ³ (STP) / h) gas mixture by combining two gases by means of a static mixer Obtained from the mixture.

The first gas mixture has a temperature of 500 ° C and a pressure of 2.46 bar and comprises: 1.06 m ³ (STP) / h of crude propane comprising propane (fresh propane) up to > 98 vol% (T = 237 ° C; P = 2.90 bar), and 0.41 m ³ (STP) / h of compressed air (T = 159 ° C; P = 2.66 bar).

The indirect heat exchange with the evolved gas A in the direction of the downstream blast furnace reactor (in the intercooler, the heat exchanger is placed between the two reactors) produces a temperature of 500 °C.

The second gas mixture comprises: 3.48 m ³ (STP) / h of dehydrogenation cycle gas (T = 554 ° C; P = 2.37 bar) and 1.98 m ³ (STP) / h of propane cycle gas (T = 500 ° C; P = 2.97 bar).

It has a pressure of 2.46 bar and a temperature of 529 °C.

The heat exchange with the product gas A leading to the downstream blast furnace reactor produces a temperature of the propane recycle gas. The dehydrogenation cycle gas stream is achieved according to the principle of a jet pump (cf. DE-A 10211275), and the propane cycle gas heated to 500 ° C acts as an actuating jet.

Adjusting the inlet temperature of the reaction gas A charging gas mixture and the bed height of the first catalyst bed through which the gas mixture is filled by the reaction gas A in the flow direction in the disk reactor so that the reaction gas A having the following contents is 522 The temperature of ° C and the pressure of 2.44 bar left the catalyst bed: propane 53.77 vol%, propylene 14.45 vol%, ethylene 0.29 vol%, methane 1.26 vol%, H ₂ 4.31 vol%, O ₂ 0 vol%, N ₂ 15.91 vol% , H ₂ O 8.88 vol%, CO 0 vol%, CO ₂ 0.01 vol%.

The amount of departure is 7.18 m ³ (STP) / h. When the first catalyst bed of the disc reactor was passed, 0.64 m ³ (STP) / h of air was introduced into the reaction gas A (T = 159 ° C, P = 2.66 bar). The metered addition was achieved by throttling so that the resulting pressure of the resulting reaction gas A was still 2.44 bar.

The bed height of the second catalyst bed of the tray reactor was such that the reaction gas A having the following contents left the second catalyst bed at a temperature of 539 ° C and a pressure of 2.41 bar: propane 43.71 vol%, propylene 17.02 vol%, ethylene 0.51% by volume, methane 1.52% by volume, H ₂ 4.53% by volume, O ₂ 0% by volume, N ₂ 20.35 vol%, H ₂ O 11.37 vol%, CO 0 vol%, and CO ₂ 0.01% by volume.

The amount of departure is 8.04 m ³ (STP) / h.

Upstream of the static mixer placed upstream of the third catalyst bed of the tray reactor, 0.63 m ³ (STP) / h of air was introduced into the reaction gas A (T = 159 ° C, at a pressure of 2.66 bar) Lower throttle). The resulting pressure of the obtained reaction gas A was still 2.41 bar.

The bed height of the third catalyst bed of the tray reactor was such that the reaction gas A having the following contents left the third catalyst bed at a temperature of 554 ° C and a pressure of 2.37 bar: propane 35.58 vol%, propylene 19.15 vol%, ethylene 0.57 vol%, methane 1.72 vol%, H ₂ 4.82 vol%, O ₂ 0 vol%, N ₂ 23.83 vol%, H ₂ O 13.31 vol%, CO 0 vol% and CO ₂ 0.01 vol%.

The amount of departure is 8.89 m ³ (STP) / h.

The splitter divides the reaction gas A into two portions having the same composition at the outlet of the disc reactor. The first part is 3.49 m ³ (STP)/h and the second part is 5.40 m ³ (STP)/h. The first portion is recycled to the reaction zone A as a component of the reaction gas A charging gas mixture. For the purpose of the first part of the dehydrogenation recycle gas stream, the heated propane recycle gas is used as an actuating injection to operate the jet pump, and the transport direction of the decelerated actuated jet passes through the mixing zone through the actuating nozzle, And the diffuser is directed in the direction of the inlet of the reaction zone A, and the nozzle is directed in the direction of the first flow portion of the reaction gas A, and the "nozzle-mixing zone-diffuser" is connected to form the first part of the reaction gas A to be recycled The only connecting line between the picking up of reaction zone A.

The second portion of the reaction gas A is led to the reaction zone A and first cooled from 554 ° C to 473 ° C by a first indirect heat exchange with the first gas mixture, the first gas mixture comprising crude propane and molecular oxygen, included in The reaction gas A starting mixture (filling gas mixture) is placed in an intercooler between the disc reactor and its downstream of the blast furnace reactor. This heats the first gas mixture from 229 ° C to 500 ° C; essentially maintaining the working pressure.

The air of 6406 m ³ (STP) / h is then metered into the second portion of the reaction gas A as described (T = 159 ° C, throttling to 2.66 bar). The gas mixture is then directed through a downstream blast furnace reactor.

The bed height of the ("fourth") catalyst bed disposed in the downstream blast furnace reactor is such that the reaction gas A exits the catalyst bed as a product gas A having the following contents at a temperature of 580 ° C and a pressure of 2.26 bar: propane 32.53 vol%, propylene 17.48 vol%, ethylene 0.52 vol%, methane 1.57 vol%, H ₂ 0 vol%, O ₂ 0 vol%, N ₂ 30.08 vol%, H ₂ O 16.79 vol%, CO 0 vol%, and CO ₂ 0.01% by volume.

In the downstream (second) indirect heat exchange with the propane recycle gas from separation zone IV, product gas A (5.92 m ³ (STP) / h) was cooled from 580 ° C to 388 ° C. Conversely, the temperature of the propane recycle gas increased from 69 ° C to 500 ° C.

Subsequently, 0.82 m ³ (STP) / h of steam was removed by condensation of product gas A from cooling to 388 ° C. To this end, the product gas A, which is cooled to 388 ° C, is first and the appropriate amount of steam has been previously removed by condensation. 5.10 m ³ (STP) / h of product gas A ^* (T = 40 ° C; P = 2.01 bar) Indirect heat exchange, and therefore cooled to 112 °C. The temperature of the product gas A ^* was simultaneously increased to 358 °C. In the downstream air cooler, the product gas A (T = 1212 ° C; P = 2.26 bar) is cooled to 54 ° C by heat exchange with external air (the working pressure of the product gas A simultaneously drops to 2.06 bar) and It has been biphasic.

Direct cooling of the atomized aqueous condensate that has been cooled (to 29 ° C) and pre-removed from product gas A by suitable direct cooling removes the above amount of steam from product gas A by condensation to obtain product gas A. ^* (P = 2.01 bar, T = 40 ° C, 5.10 m ³ (STP) / h). It was heated to 358 ° C as described above. 8.55 m ³ (STP) / h of air (T = 159 ° C, P = 2.66 bar) was then throttled into the product gas A ^* . It forms a reaction gas B starting mixture (filling gas mixture) having a temperature of 283 ° C and an operating pressure of 1.96 bar.

The starting mixture of reaction gas B (filling gas mixture) had the following contents: propane 14.11% by volume, propylene 7.58 vol%, ethylene 0.23 vol%, methane 0.68 vol%, H ₂ 0 vol%, O ₂ 12.69 vol%, N ₂ 60.9 vol%, H ₂ O 2.75 vol%, CO 0 vol%, and CO ₂ 0.03 vol%.

It is used to load the reaction zone B having the same structure as that of the structure from II.

The propylene loading of the catalyst packing in the first oxidation stage is selected to be 130 l(STP)/l. h. The salt melt (53 wt% KNO ₃ , 40 wt% NaNO ₂ , 7 wt% NaNO ₃ ) has the following inlet temperatures: T _A = 327 ° C T _B = 329 ° C T _C = 267 ° C T _D = 268 ° C .

A sufficient amount of air (159 ° C, 2.66 bar) was metered (throttled) into the product gas mixture of the first oxidation stage, and the O ₂ : acrolein molar ratio in the charge gas mixture of the second oxidation stage was 1.01.

The acrolein loading of the catalyst packing in the second oxidation stage is 110 l(STP)/l. h. The pressure at the inlet of the second oxidation stage was 1.68 bar.

The reaction gas B exited the intercooler at a temperature of 260 ° C and the inlet temperature of the charged gas mixture into the second oxidation stage was 253 ° C.

The product gas mixture of the first oxidation stage has the following contents: 6.40% by volume of acrolein, 0.44% by volume of acrylic acid, 14.09% by volume of propane, 0.39% by volume of propylene, 0.68% by volume of methane, H ₂ 0% by volume, and 0.23% by volume of ethylene. O ₂ 4.18 vol%, N ₂ 60.87 vol%, H ₂ O 10.64 vol%, CO 0.33 vol%, and CO ₂ 0.78 vol%.

The temperature of the product gas in the first oxidation stage was 335 ° C before entering the aftercooler.

The product gas mixture of the second oxidation stage (product gas B) has a temperature of 270 ° C and a pressure of 1.55 bar, and the following contents: acrolein 0.03 vol%, acrylic acid 5.97 vol%, acetic acid 0.18 vol%, propane 13.0 vol% , propylene 0.36 vol%, methane 0.63 vol%, ethylene 0.21 vol%, H ₂ 0 vol%, O ₂ 2.60 vol%, N ₂ 64.2 vol%, H ₂ O 10.29 vol%, CO 0.48 vol% and CO ₂ 1.08 vol %.

Product gas B (14.81 m ³ (STP) / h) was fractionally condensed in a tray column (separation zone II) as described in WO 2004/035514.

As the first fuel, 13.8 g/h of a high boiling point (polyacrylic acid (Michael adduct) or the like) was sent to the residue for incineration.

From the second collection tray above the feed of product gas B in the tray column, 2.80 kg/h of condensed crude acrylic acid having a temperature of 15 ° C and 96.99 wt% of acrylic acid was withdrawn. As described in WO 2004/035514, it is suspended and crystallized after the addition of a small amount of water, the suspended crystals are separated from the mother liquor in a hydraulic scrubber and the mother liquor is recycled to the condensation column as described in WO 2004/035514. The washed suspension crystals have a purity of >99.87 wt% acrylic acid and are suitable for the immediate preparation of polymers for "superabsorbent" water in sanitary articles.

The amount of acid water condensate drawn from the third collection tray above the feed to the product gas B of the condensation column but not recycled to the condensation column is 1.19 kg/h, has a temperature of 34 ° C and has the following inclusions : 0.35 wt% acrolein, 4.95 wt% acrylic acid, 5.31 wt% acetic acid, and 88.40 wt% water.

At the top of the condensation column, the residual gas of 12.56 m ³ (STP) / h leaves the separation zone II at a temperature of 33 ° C and a pressure of 1.20 bar and below: acryl 0.02 vol%, acrylic acid 0.02 vol%, acetic acid 0 % by volume, propane 15.35 vol%, propylene 0.42 vol%, ethylene 0.25 vol%, methane 0.74 vol%, H ₂ 0 vol%, O ₂ 3.07 vol%, N ₂ 75.64 vol%, H ₂ O 1.68 vol%, CO 0.55 % by volume, and 1.27% by volume of CO ₂ .

In the first compressor stage of the multistage radial compressor, the residual gas is compressed from 1.20 bar to 6.93 bar, during which the temperature of the residual gas rises to 217 °C.

In the indirect heat exchanger, the residual gas is cooled to 111 ° C without condensate formation (the coolant is "propane washed" residual gas (which is also referred to as "exhaust gas" hereinafter), in the indirect heat exchanger from 30 After heating to 138 ° C, °C begins to decompress from 20 bar in the first expansion stage of the multistage expansion turbine and simultaneously cools). This heated exhaust gas (10.39 m ³ (STP) / h). In the second expansion stage, the exhaust gas is depressurized to 1.10 bar (which subsequently has a temperature of 78 ° C) and then sent to the residue for incineration.

The compressed residual gas (P = 6.93 bar and T = 1111 ° C) was cooled to 59 ° C in an indirect air cooler. In the second compressor stage, it is compressed from 6.93 bar to 20 bar while it is heated to 168 °C. In the downstream indirect heat exchanger, it is cooled to 101 ° C (the coolant is "propane washed" residual gas, which is simultaneously heated from 30 ° C to 138 ° C) without condensate formation. In the downstream of another indirect heat exchanger, the residual gas at 101 ° C was cooled to 35 ° C (the coolant was surface water of about 25 ° C).

It condenses 159.7 g/h of water from the residual gas. The aqueous condensate is likewise sent to the residue for incineration and is incinerated therein along with the other residues mentioned while supplying 5.59 m ³ STP) / h of air. The hot combustion gases were cooled to obtain steam (52 bar, 267 ° C, 3.78 kg/h) and released into the environment.

The residual gas, which still contains 12.37 m ³ (STP) / h of propane (20 bar, 35 ° C), is introduced into the lower section of the column (separation zone III) with the random packing.

At the top of the scrubber, 66.37 kg/h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 °C (P = 20 bar).

At the top of the scrubber, the "propane-washed" residual gas (exhaust gas) escapes at a pressure of 20 bar and a temperature of 30 ° C and an amount of 10.39 m ³ (STP) / h, which has the following inclusions: Propane 0.19 vol%, propylene 0.01 vol%, ethylene 0.29 vol%, methane 0.89 vol%, H ₂ 0 vol%, O ₂ 3.70 vol%, N ₂ 91.38 vol%, H ₂ O 0.32 vol%, CO 0.67 vol%, And CO ₂ 1.53 vol%.

The exhaust gas is subjected to indirect heat exchange as described and then sent to the residue for incineration or by means of flame incineration.

70.25 kg/h of absorbent having a pressure of 20 bar and a temperature of 41 ° C was withdrawn from the bottom of the "propane scrubber" which was not externally cooled or heated. It comprises 5.41% by weight of propane and 0.14% by weight of propylene.

It is heated to 69 ° C in an indirect heat exchanger before the absorber is directed to the top of the downstream desorption column. The heat carrier used was a liquid effluent from the desorption column (66.37 kg/h) which had a temperature of 80 ° C at a pressure of 2.97 bar and still included 0.03 wt% of dissolved propane. It was compressed back to 20 bar by means of a pump and after being indirectly cooled to 30 ° C with surface water (25 ° C), it was recycled to the top of the "propane scrubber" to propane absorption.

After heating to 69 ° C, the absorbent (for example in a reverse pump or by means of a valve) is decompressed to a pressure of 2.97 bar (the mechanical energy released under reverse pump conditions is also used to recompress the desorption tower) A propane-containing absorbent (liquid effluent from the desorption column)), and the obtained two-phase mixture is introduced into the desorption column at the top.

Disposed in the desorption column is a heating coil that directs steam having a temperature of 152 ° C, P = 5.00 bar, and 1.3 kg / h through the heating coil to countercurrent the descending absorbent in the desorption column.

At the top of the desorption column, a propane recycle gas of 1.98 m ³ (STP)/h having a temperature of 69 ° C and a working pressure of 2.97 bar was released. After heating to 500 ° C by heat exchange with the product gas A having a temperature of T = 580 ° C, the propane cycle gas is used as an actuating injection to operate the jet pump, and the dehydrogenation cycle gas and the propane recycle gas system are pumped by the jet Filling to reaction zone A using the starting mixture of reaction gas A.

The propane recycle gas had the following contents: propane 96.42% by volume and propylene 2.58% by volume.

VI. First comparative example of a method for preparing acrylic acid from propane (describes stable operating state)

Reaction zone A consists of an adiabatic blast furnace reactor having only one fixed catalyst bed, the configuration of which corresponds to the first fixed bed in the flow direction from the disc reactor of II.

The heterogeneously catalyzed partial propane dehydrogenation is carried out in a blast furnace reactor in a direct transfer mode with a dehydrogenation recycle gas. The total amount of catalyst on the propane loading (calculated without inert material) is 1500 1 (STP) / 1. h.

24.87 m ³ (STP) / h of the reaction gas A filling gas mixture (starting mixture: T = 450 ° C; P = 2.82 bar) was sent to the blast furnace reactor and had the following contents: propane 22.3 vol%, propylene 3.65 vol %, ethane 0.07 vol%, methane 0.17 vol%, H ₂ 4.18 vol%, O ₂ 2.09 vol%, N ₂ 60.32 vol%, H ₂ O 5.05 vol%, CO 0 vol%, and CO ₂ 1.17 vol%.

The reaction gas A is filled with a gas mixture as a mixture of a first (1.69 m ³ (STP) / h) and a second (23.18 m ³ (STP) / h) gas mixture by combining two gases by means of a static mixer Obtained from the mixture.

The first gas mixture is 0.013 m ³ (STP) / h of steam (T = 134 ° C; P = 3 bar), 1.06 m ³ (STP) / h including propane (fresh propane) to the extent of > 98 vol% a mixture of crude propane (T = 122 ° C; P = 3.50 bar) and 0.62 m ³ (STP) / h of molecular hydrogen (purity > 99% by volume; T = 35 ° C; P = 2.20 bar), and a second gas The mixture was a compressed residual gas cycle gas of 11.41 m ³ (STP) / h of dehydrogenation cycle gas (T = 534 ° C; P = 2.36 bar), 11.16 m ³ (STP) / h (T = 444 ° C; P = 2.82 bar) and a mixture of compressed air (T = 175 ° C, P = 3.00 bar) of 0.61 m ³ (STP) / h.

Indirect heat exchange with the derivatized reaction zone A and product gas A in the direction of reaction zone B produces a temperature of the compressed residual gas recycle gas at 444 °C.

The dehydrogenation cycle gas stream is achieved by the jet pump principle (cf. DE-A 102 112 75), which sets the compressed residual gas recycle gas to 444 ° C to actuate the injection. The compressed air is metered (throttled) into the resulting mixture.

The bed height of the fixed catalyst bed through which the gas mixture was charged by the reaction gas A was adjusted so that the product gas A having the following contents left the catalyst bed at a temperature of 534 ° C and a pressure of 2.36 bar: propane 17.83 vol%, propylene 7.51 vol. %, ethane 0.15 vol%, methane 0.37 vol%, H ₂ 3.76 vol%, O ₂ 0 vol%, N ₂ 59.18 vol%, H ₂ O 9.06 vol%, CO 0 vol%, and CO ₂ 1.14 vol%.

The amount of departure is 25.35 m ³ (STP) / h.

The splitter divides the product gas A at the outlet of the blast furnace reactor into two fractions having the same composition.

The first part is 11.41 m ³ (STP)/h and the second part is 13.94 m ³ (STP)/h. The first portion is recycled to the reaction zone A as a component of the reaction gas A charging gas mixture. For the purpose of the first part of the dehydrogenation recycle gas stream, the jet pump is operated by the heated compressed residual gas recycle gas as the actuating injection, and the transport direction of the actuated jet through the decompression is passed through the mixing zone. Actuating the nozzle, and the diffuser is directed in the direction of the inlet into the reaction zone A, and the nozzle is directed in the direction of the first flow portion of the product gas A, and the "nozzle-mixing zone-diffuser" connection is formed to be recirculated The only connecting line between the first portion of product gas A and the reaction zone A.

The second portion of product gas A is led to reaction zone A and placed in the downstream heat exchanger between it and by the compressed residual gas cycle gas (11.16 m ³ (STP) / h; T = 131 ° C; P = 3.00 bar) was exchanged for heat exchange to 317 ° C (basically maintaining working pressure). This heats the residual gas recycle gas to 444 ° C and reduces its working pressure to 2.82 bar.

Subsequently, the product gas A is cooled to 221 ° C in a second indirect heat exchange with the "propane-washed" product gas A (9.06 m ³ (STP) / h; T = 30 ° C, P = 10.50 bar), accompanied by Its working pressure is reduced to 2.18 bar.

This heats the "propane washed" product gas A (hereinafter also referred to as "exhaust gas") to 287 ° C and changes its working pressure to 10.10 bar. Subsequently, the exhaust gas is depressurized from 10.10 bar to 3.50 bar in the first expansion stage of the multistage expansion turbine, and simultaneously cooled from 287 ° C to 176 ° C. The exhaust gas leaving the first expansion stage is then indirectly heat exchanged with the product gas A at 221 °C. The latter essentially maintained its working pressure and simultaneously cooled to 215 °C. It heats the exhaust gases to 191 ° C while maintaining their working pressure substantially. The offgas is then depressurized to 1.23 bar in the second expansion stage of the multistage expansion turbine, which is allowed to cool to 89 °C. The exhaust gas is then sent to the residue for incineration.

The product gas A at 215 ° C and 2.18 bar was then first cooled to 60 ° C in an air cooled inter-exchange heat exchanger and then cooled to 39 ° C in an intercooled heat exchanger cooled by cooling water. The water condensed during the cooling process (0.68 kg/h) was removed by means of a small droplet separator.

The remaining 13.09 m ³ (STP) / h of product gas A was compressed to 4.78 bar in the first compressor stage of the multistage radial compressor. This heats product gas A from 39 ° C to 106 ° C. Subsequently, the product gas A was cooled in air to 54 ° C in an indirect heat exchanger and the condensed water (8.0 g/h) was removed by means of a small droplet separator. In the second compression stage, the remaining 13.08 m ³ (STP) / h of product gas A was compressed to 10.50 bar. It is related to heating of product gas A to 123 °C. Subsequently, the product gas A was first cooled to 54 ° C by cooling with air in an indirect heat exchanger and then cooled to 29 ° C by heat exchange with cooling water. The water condensed during the cooling process (0.29 kg/h) was removed by means of a small droplet separator. The still remaining product gas A of 12.71 m ³ (STP) / h (whose operating pressure is still substantially 10.50 bar) is introduced into the lower portion of the column with the random packing. The contents are: propane 19.55 vol%, propylene 8.24 vol%, ethane 0.16 vol%, methane 0.41 vol%, H ₂ 4.13 vol%, O ₂ vol%, N ₂ 64.9 vol%, H ₂ O 0.38 vol %, CO 0% by volume, and CO ₂ 1.25 vol%.

At the top of the propane scrubber, 120.32 kg/h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 °C (P = 10.50 bar).

At the top of the scrubber, the "propane-washed" product gas A (exhaust gas) escapes, with a pressure of 10.50 bar and a temperature of 30 ° C and a volume of 9.06 m ³ (STP) / h, which has the following inclusions : propane 0.02 vol%, propylene 0.01 vol%, ethane 0.14 vol%, methane 0.56 vol%, H ₂ 5.79 vol%, O ₂ 0 vol%, N ₂ 91.0 vol%, H ₂ O 0.40 vol%, CO 0 volume %, and CO ₂ 1.07% by volume.

After heat exchange is carried out as described above with the product gas A from the derivation reaction zone A and multistage depressurization to 1.23 bar in a multistage expansion turbine, the offgas is sent to the residue for incineration or incineration by flame. It may be appropriate to remove the molecular hydrogen present therein on an industrial scale before the exhaust gas is decompressed. This can be achieved, for example, by separating the exhaust gases through the membrane as already described. The removed molecular hydrogen can thus be completely or partially recycled to reaction zone A (suitably as component of the reaction gas A charge gas mixture). Alternatively, it can be burned, for example, in a fuel unit.

From the bottom of the "propane scrubber" which was not externally cooled or heated, 127.37 kg/h of absorbent having a pressure of 10.5 bar and a temperature of 40 ° C was withdrawn. Which comprises 3.83% by weight of propane, 1.54% by weight of propylene, 0.1% by weight of ethane and 0.1% by weight of CO _2. Before the absorber is directed to the top of the downstream desorption column, it is depressurized to a pressure of 2.01 bar (for example in a reverse pump or by means of a valve; the mechanical energy released in the case of a reverse pump is also suitable for recompressing the desorption tower Propane-free absorbent (liquid discharge from the desorption column)). The two-phase mixture obtained is introduced into the desorption column at the top of the desorption column.

The air of 8.00 m ³ (STP)/h compressed to a pressure of 3.00 bar (which is heated to 175 ° C) is introduced as a stripping gas from the bottom and the absorbent falling from the top of the desorption column into the desorption column. Also has a tower of random packing).

The liquid discharge (120.32 kg/h) of the desorption column was compressed back to 10.50 bar by means of a pump and, after indirect cooling with cooling water, was recycled to the top of the propane scrubber (absorption column) at a temperature of 30 °C. The propane is absorbed.

At the top of the desorption column, a gas mixture of 11.64 m ³ (STP) / h was allowed to escape at a temperature of 33 ° C and a pressure of 2.01 bar and with the following contents: propane 21.33 vol%, propylene 8.99 vol%, ethane 0.07 vol%, O ₂ 13.93 vol%, N ₂ 52.6 vol%, H ₂ O 1.53 vol%, CO 0 vol%, and CO ₂ 0.55 vol%.

The heat exchange with the steam increases the temperature of the gas mixture to 140 ° C (this reduces the working pressure to 1.91 bar).

This gas mixture was used as a starting mixture of reaction gas B (filling gas mixture) to charge reaction zone B having the same structure as that of structure II.

The propylene loading of the catalyst packing in the first oxidation stage was selected to be 130 l (STP) / h. The salt melt (53 wt% KNO ₃ , 40 wt% NaNO ₂ , 7 wt% NaNO ₃ ) has the following inlet temperatures: T _A = 329 ° C T _B = 330 ° C T _C = 265 ° C T _D = 266 ° C .

The reaction gas B leaves the intercooler at a temperature of 260 ° C and a pressure of 1.68 bar and below at a pressure of 11.66 m ³ (STP) / h: 7.51 vol% of acrolein, 0.51 vol% of acrylic acid, and 21.3% by volume of propane. Propylene 0.53 vol%, ethane 0.07 vol%, O ₂ 3.92 vol%, N ₂ 52.52 vol%, H ₂ O 10.81 vol%, CO 0.38 vol%, and CO ₂ 1.43 vol%.

The air of 2.17 m ³ (STP)/h compressed in a radial compressor to 3 bar (which is heated to 175 ° C) is metered (throttled) into the reaction gas at an inlet pressure of 1.68 bar. The inlet temperature to the second oxidation stage was 252 ° C (acrolein loading = 1010 l (STP) / l.h).

The product gas mixture of the second oxidation stage (product gas B) has a temperature of 270 ° C and a pressure of 1.55 bar and the following contents: acrolein 0.03 vol%, acrylic acid 6.62 vol%, acetic acid 0.2 vol%, propane 18.51 vol%, Propylene 0.46 vol%, ethane 0.06 vol%, O ₂ 2.95 vol%, N ₂ 58.0 vol%, H ₂ O 10.0 vol%, CO 0.52 vol%, and CO ₂ 1.64 vol%.

As described in WO 2004/035514, product gas B (13.42 m ³ (STP)/h) is fractionally condensed in a tray column (separation zone II).

A high boiling point of 13.8 g/h (polyacrylic acid (Michael adduct), etc.) was sent to the residue for incineration.

From the second collection tray above the feed of product gas B in the tray column, 2.82 kg/h of condensed crude acrylic acid having a temperature of 15 ° C and 96.99 wt% of acrylic acid was withdrawn. As described in WO 2004/035514, it is suspended and crystallized after the addition of a small amount of water, the suspended crystals are separated from the mother liquor in a hydraulic scrubber and the mother liquor is recycled to the condensation column as described in WO 2004/035514. The washed suspension crystals have a purity of >99.87 wt% acrylic acid and are suitable for the immediate preparation of polymers for "superabsorbent" water in sanitary articles.

The amount of acid water condensate withdrawn from the third collection tray above the feed of product gas B entering the condensation column and not recycled to the condensation column was 1.04 kg/h and had a temperature of 33 °C. It is also sent to the residue for incineration.

At the top of the condensation column, 11.16 m ³ (STP) / h of residual gas leaves the condensation column at a temperature of 33 ° C and a pressure of 1.20 bar and below: propane 22.06 vol%, propylene 0.46 vol%, O ₂ 3.54 volume %, N ₂ 69.75 vol%, H ₂ O 1.76 vol%, CO 0 vol% and CO ₂ 1.43 vol%.

The residual gas is all compressed to 3 bar (this is heated to 131 ° C) and, as described, after the heat exchange with the product gas A of the derivation reaction zone A, as a residual gas recycle gas (as a dehydrogenation cycle) The actuating injection of the jet of the gas stream is recycled to the reaction zone A.

VII. A second comparative example of a method for preparing acrylic acid from propane (describes a stable operating state)

Reaction zone A consists of an adiabatic blast furnace reactor having only one fixed catalyst bed, the configuration of which corresponds to the first fixed bed in the flow direction from the disc reactor of II.

The heterogeneously catalyzed partial propane dehydrogenation is carried out in a blast furnace reactor in a direct transfer mode with a dehydrogenation recycle gas. The total amount of catalyst on the propane loading (calculated without inert material) is 1500 1 (STP) / 1. h.

23.79 m ³ (STP) / h of the reaction gas A filling gas mixture (starting mixture: T = 456 ° C; P = 2.82 bar) was sent to the blast furnace reactor and had the following contents: propane 19.31 vol%, propylene 3.82 vol %, ethane 0.11% by volume, methane 0.33% by volume, H ₂ 4.27 vol%, O ₂ 2.14 vol%, N ₂ 62.72 vol%, H ₂ O 5.11 vol%, CO 0 vol%, and CO ₂ 1.22 vol%.

The reaction gas A is filled with a gas mixture as a mixture of a first (1.69 m ³ (STP) / h) and a second (22.10 m ³ (STP) / h) gas mixture by combining two gases by means of a static mixer Obtained from the mixture.

The first gas mixture is 0.013 m ³ (STP) / h of steam (T = 134 ° C; P = 3 bar), 1.08 m ³ (STP) / h including propane (fresh propane) to the extent of > 98 vol% a mixture of crude propane (T = 122 ° C; P = 3.50 bar) and 0.60 m ³ (STP) / h of molecular hydrogen (purity > 99% by volume; T = 35 ° C; P = 2.20 bar), and a second gas The mixture was a dehydrogenation cycle gas of 10.93 m ³ (STP) / h (T = 546 ° C; P = 2.82 bar), and a compressed residual gas cycle gas of 10.52 m ³ (STP) / h (P = 3 bar; T = A mixture of 455 ° C) and 0.65 m ³ (STP) / h of compressed air (T = 175 ° C, P = 3.00 bar).

Indirect heat exchange with the derivatized reaction zone A and product gas A in the direction of reaction zone B produces a temperature of the compressed residual gas recycle gas at 455 °C.

The dehydrogenation cycle gas stream is achieved by the jet pump principle (cf. DE-A 102 112 75), which sets the compressed residual gas recycle gas to 455 ° C to actuate the injection. The compressed air is metered (throttled) into the resulting mixture.

The bed height of the fixed catalyst bed through which the gas mixture was charged by the reaction gas A was such that the product gas A having the following contents left the catalyst bed at a temperature of 546 ° C and a pressure of 2.39 bar: propane 14.60 vol%, propylene 7.84 vol. %, ethane 0.24% by volume, methane 0.70% by volume, H ₂ 3.79 vol%, O ₂ 0 vol%, N ₂ 61.44 vol%, H ₂ O 9.19 vol%, CO 0 vol%, and CO ₂ 1.19 vol%.

The amount of departure is 24.29 m ³ (STP) / h.

The splitter divides the product gas A at the outlet of the blast furnace reactor into two fractions having the same composition.

The first part is 10.93 m ³ (STP)/h and the second part is 13.36 m ³ (STP)/h. The first portion is recycled to the reaction zone A as a component of the reaction gas A charging gas mixture. For the purpose of the first part of the dehydrogenation recycle gas stream, the jet pump is operated by the heated compressed residual gas recycle gas as the actuating injection, and the transport direction of the actuated jet through the decompression is passed through the mixing zone. Actuating the nozzle, and the diffuser is directed in the direction of the inlet into the reaction zone A, and the nozzle is directed in the direction of the first flow portion of the product gas A, and the "nozzle-mixing zone-diffuser" connection is formed to be recirculated The only connecting line between the first portion of product gas A and the reaction zone A.

The second portion of product gas A is directed to reaction zone A and placed in the downstream heat exchanger between it and by the compressed residual gas cycle gas (10.52 m ³ (STP) / h; T = 135 ° C; P = 3.00 bar) was exchanged for heat exchange to 317 ° C (basically maintaining working pressure). This heats the residual gas recycle gas to 455 ° C and reduces its working pressure to 2.82 bar.

Subsequently, the product gas A is cooled to 220 ° C in a second indirect heat exchange with the "propane-washed" product gas A (9.06 m ³ (STP) / h; T = 30 ° C, P = 10.50 bar), accompanied by Its working pressure is reduced to 2.21 bar.

This heats the "propane washed" product gas A (hereinafter also referred to as "exhaust gas") to 300 ° C and changes its working pressure to 10.10 bar. Subsequently, the exhaust gas was depressurized from 10.10 bar to 3.50 bar in the first expansion stage of the multistage expansion turbine, and simultaneously cooled from 300 ° C to 187 ° C. The off-gas from the first expansion stage is then directed to indirect heat exchange with the product gas A at 220 °C. The latter essentially maintained its working pressure and simultaneously cooled to 219 °C. It heats the exhaust gases to 190 ° C while maintaining their working pressure substantially. The exhaust gas is then depressurized to 1.23 bar in the second expansion stage of the multistage expansion turbine, which is allowed to cool to 88 °C. The exhaust gas is then sent to the residue for incineration.

The product gas A at 219 ° C and 2.21 bar was then first cooled to 60 ° C in an air cooled inter-exchange heat exchanger and then cooled to 39 ° C in an intercooled heat exchanger cooled by cooling water. The water condensed during the cooling process (0.68 kg/h) was removed by means of a small droplet separator.

The remaining 12.52 m ³ (STP) / h of product gas A was compressed to 4.81 bar in the first compressor stage of the multistage radial compressor. This heats the product gas A from 39 ° C to 108 ° C; subsequently, the product gas A is cooled in air to 54 ° C in an indirect heat exchanger and the condensed water is removed by means of a small droplet separator (5.7 g / h ). In the second compression stage, the remaining 12.51 m ³ (STP) / h of product gas A was compressed to 10.50 bar. This is related to heating product gas A to 126 °C. Subsequently, the product gas A was first cooled to 54 ° C by cooling with air in an indirect heat exchanger and then cooled to 29 ° C by heat exchange with cooling water. The water condensed during the cooling process (0.28 kg/h) was removed by means of a small droplet separator. The still remaining product gas A of 12.17 m ³ (STP) / h (whose operating pressure is still substantially 10.50 bar) is introduced into the lower portion of the column with the random packing. The contents are: propane 16.0 vol%, propylene 8.60 vol%, ethane 0.26 vol%, methane 0.77 vol%, H ₂ 4.17 vol%, O ₂ 0 vol%, N ₂ 67.48 vol%, H ₂ O 0.39 vol %, CO 0% by volume, and CO ₂ 1.31% by volume.

At the top of the propane scrubber, 115.96 kg/h of industrial tetradecane from Haltermann, Germany (as described in II) was introduced at an introduction temperature of 30 °C (P = 10.50 bar).

At the top of the scrubber, the "propane-washed" product gas A (exhaust gas) escapes, with a pressure of 10.50 bar and a temperature of 30 ° C and a volume of 9.06 m ³ (STP) / h, which has the following inclusions : propane 0.02% by volume, propylene 0.01% by volume, ethane 0.22% by volume, methane 1.04% by volume, H ₂ 5.59 vol%, O ₂ 0 vol%, N ₂ 90.60 vol%, H ₂ O 0.39 vol%, CO 0 volume %, and CO ₂ 1.12% by volume.

After heat exchange is carried out as described above with the product gas A of the derivation reaction zone A and multistage depressurization to 1.23 bar in a multistage expansion turbine, the exhaust gas is sent to the residue for incineration. It may be appropriate to remove the molecular hydrogen present therein on an industrial scale before the exhaust gas is decompressed. This can be achieved, for example, by separating the exhaust gases through the membrane as already described. The removed molecular hydrogen can thus be completely or partially recycled to reaction zone A (suitably as component of the reaction gas A charge gas mixture). Alternatively, it can be burned, for example, in a fuel unit.

From the bottom of the "propane scrubber" which was not externally cooled or heated, 121.93 kg/h of the absorbent having a pressure of 10.5 bar and a temperature of 40 ° C was withdrawn. It comprises 3.14% by weight of propane, 1.61% by weight of propylene, 0.01% by weight of ethane and 0.09% by weight of CO ₂ . Before the absorber is directed to the top of the downstream desorption column, it is depressurized to a pressure of 1.98 bar (for example in a reverse pump or by means of a valve; the mechanical energy released in the case of a reverse pump is also suitable for recompressing the desorption tower Propane-free absorbent (liquid discharge from the desorption column)). The two-phase mixture obtained is introduced into the desorption column at the top of the desorption column.

7.88 m ³ (STP) / h of air compressed to a pressure of 3.00 bar (which is heated to 175 ° C) as a stripping gas from the bottom and the absorber falling from the top of the desorption tower is introduced into the desorption column countercurrently ( Also has a tower of random packing).

The liquid discharge (115.94 kg/h) of the desorption column was compressed back to 10.50 bar by means of a pump and, after indirect cooling with cooling water, was recycled to the top of the propane scrubber (absorption column) at a temperature of 30 °C. The propane is absorbed.

At the top of the desorption column, a gas mixture of 10.97 m ³ (STP) / h was allowed to escape at a temperature of 33 ° C and a pressure of 1.98 bar and with the following contents: propane 17.77 vol%, propylene 9.53 vol%, ethane 0.11% by volume, O ₂ 14.54% by volume, N ₂ 54.91% by volume, H ₂ O 1.59% by volume, CO 0% by volume, and CO ₂ 0.55% by volume.

The heat exchange with the steam increases the temperature of the gas mixture to 140 ° C (this reduces the working pressure to 1.88 bar).

This gas mixture was used as a starting mixture of reaction gas B (filling gas mixture) to charge reaction zone B having the same structure as that of structure II.

The propylene loading of the catalyst packing in the first oxidation stage was selected to be 130 l (STP) / h. Salt melt (53% by weight of KNO _3, 40 wt% of NaNO _2, 7% by weight of NaNO ₃₎ has an inlet _{temperature: T A = 323 ℃ T B} = 327 ℃ T C = 262 ℃ T D = 265. C.

B in an amount of reaction gas (STP) / h of 10.99 m ³ and a temperature of 260 ℃, pressure of 1.67 bar and the contents of exits intercooler: 7.98% by volume acrolein, 0.53 volume percent acrylic acid, 17.74 vol% of propane, Propylene 0.56 vol%, ethane 0.1 vol%, O ₂ 3.92 vol%, N ₂ 54.82 vol%, H ₂ O 11.44 vol%, CO 0.41 vol%, and CO ₂ 1.49 vol%.

The air of 2.21 m ³ (STP) / h compressed in a radial compressor to 3 bar (which is heated to 175 ° C) is metered (throttled) into the reaction at an inlet pressure of 1.67 bar. The inlet temperature of the gas entering the second oxidation stage was 251 ° C (acrolein loading = 1010 l (STP) / l.h).

The product gas mixture (product gas B) of the second oxidation stage has a temperature of 270 ° C and a pressure of 1.55 bar and the following contents: acrolein 0.03 vol%, acrylic acid 6.94 vol%, acetic acid 0.21 vol%, propane 15.25 vol%, Propylene 0.48 vol%, ethane 0.09 vol%, O ₂ 2.95 vol%, N ₂ 60.31 vol%, H ₂ O 10.49 vol%, CO 0.55 vol%, and CO ₂ 1.69 vol%.

As described in WO 2004/035514, product gas B (12.79 m ³ (STP)/h) is fractionally condensed in a tray column (separation zone II).

A high boiling point of 13.8 g/h (polyacrylic acid (Michael adduct), etc.) was sent to the residue for incineration.

From the second collection tray above the feed of product gas B in the tray column, 2.82 kg/h of condensed crude acrylic acid having a temperature of 15 ° C and 96.99 wt% of acrylic acid was withdrawn. As described in WO 2004/035514, it is suspended and crystallized after the addition of a small amount of water, the suspended crystals are separated from the mother liquor in a hydraulic scrubber and the mother liquor is recycled to the condensation column as described in WO 2004/035514. The washed suspension crystals have a purity of >99.87 wt% acrylic acid and are suitable for the immediate preparation of polymers for "superabsorbent" water in sanitary articles.

The amount of acid water condensate withdrawn from the third collection tray above the feed of product gas B entering the condensation column and not recycled to the condensation column was 1.06 kg/h and had a temperature of 33 °C. It is also sent to the residue for incineration.

At the top of the condensation column, the residual gas of 10.52 m ³ (STP) / h leaves the condensation column at a temperature of 33 ° C and a pressure of 1.20 bar and below: propane 18.36 vol%, propylene 0.49 vol%, O ₂ 3.58 vol %, N ₂ 73.32 vol%, H ₂ O 1.73 vol%, CO 0 vol%, and CO ₂ 1.51 vol%.

The residual gas is all compressed to 3 bar (this is heated to 135 ° C) and, as described, after the heat exchange with the product gas A of the derivation reaction zone A, as a residual gas recycle gas (as a dehydrogenation cycle) The actuating injection of the jet of the gas stream is recycled to the reaction zone A.

U.S. Provisional Patent Application Serial No. 60/802,786, filed on May 24, 2006, is hereby incorporated by reference.

Many variations and deviations from the present invention are possible in light of the above teachings. It is therefore contemplated that the invention may be practiced otherwise than as specifically described herein within the scope of the appended claims.

Claims (33)

A process for the preparation of acrolein or acrylic acid or a mixture thereof from propane, wherein A)- a gaseous feed stream comprising at least two propanes is sent to a first reaction zone A to form a reaction gas A, wherein the at least two feed streams At least one of them comprises fresh propane; - in reaction zone A, directing reaction gas A through at least one catalyst bed, wherein a portion of the propane is heterogeneously catalytically dehydrogenated to form molecular hydrogen and propylene; - molecular oxygen And sent to the reaction zone A, and in the reaction zone A, at least a part of the molecules present in the reaction gas A are oxidized to steam, and - the product gas A including propylene, propane and steam is extracted from the reaction zone A; B) when appropriate In the first separation zone I, the vapor present in the product gas A is partially or completely condensed by the product gas A and/or direct cooling to leave the product gas A*; C) in a reaction In zone B, at least one oxidation reactor having a reaction gas B comprising propane, propylene and molecular oxygen is charged with product gas A or product gas A* and molecular oxygen feed, and the propylene present therein is heterogeneously phased. Catalytic part Gas phase oxidation to obtain acrolein or acrylic acid or a mixture thereof as a target product, and product gas B including unconverted propane; D) to export product gas B to reaction zone B, and to remove it in second separation zone II The target product present leaving a residual gas including propane E) as appropriate, a portion of the residual gas having the residual gas composition is recycled to the reaction zone A as a feed stream comprising propane; F) in a separation zone III, the propane present in the residual gas is borrowed An absorber comprising propane is formed by absorption from the residual gas into an organic solvent, wherein the residual gas has not been recycled to the reaction zone A, and any vapor present therein if appropriate has been partially or completely condensed in advance Any molecular hydrogen removed and/or present therein has been partially or completely removed by means of a separation membrane in advance; and G) in a separation zone IV, the propane is removed from the absorption and as a propane comprising The stream is recycled to the reaction zone A; the method comprises: oxidizing at least enough molecules into steam in the reaction zone A, and the amount of hydrogen oxidized into steam in the reaction zone A is the molecular hydrogen formed in the reaction zone A Amount of 30 to 70 mol%. The method of claim 1, wherein at least some of the thermal energy generated by the oxidation of the molecular hydrogen in the reaction zone A is used to react with the reaction gas A or the product gas A or the reaction gas A and the product gas A as a heat carrier. The gaseous feed stream sent to reaction zone A is heated by indirect heat exchange. The method of claim 1 or 2, wherein the amount of hydrogen oxidized to vapor in the reaction zone A is at least 40 mol% of the amount of molecular hydrogen formed in the reaction zone A. The method of claim 1 or 2, wherein the reactive gas B comprises molecular hydrogen. The method of claim 1 or 2, wherein the product gas A withdrawn from the reaction zone A is as used to charge the at least one oxidation reaction in the reaction zone B. Device. The method of claim 1 or 2, wherein at least 5 mol% of the steam present in product gas B is removed by condensation in separation zone I. The method of claim 1 or 2, wherein at least 25 mol% of the steam present in product gas B is removed by condensation in separation zone I. The method of claim 1 or 2, wherein at least 50 mol% of the steam present in product gas B is removed by condensation in separation zone I. The method of claim 1 or 2, wherein 5 to 98 mol% of the steam present in the product gas B is removed by condensation in the separation zone I. The method of claim 1 or 2, wherein the working pressure in the reaction zone A is greater than 1 bar and at most 5 bar. The method of claim 1 or 2, wherein the following relationship applies to the working pressure P in the different zones of the method, the working pressure P being measured at the entrance to the specific zone under each condition: P _{reaction zone A} > P _{Separation zone I} > P _{reaction zone B} > P _{separation zone II} < P _{separation zone III} > P _{separation zone IV} > P _{reaction zone A.} The method of claim 1 or 2, wherein the molecular hydrogen is burned together with the molecular oxygen in the reaction zone A when 90 mol% of the total amount of molecular hydrogen formed in the reaction zone A is formed. The method of claim 1 or 2, wherein the molecular hydrogen is burned together with the molecular oxygen in the reaction zone A when 75 mol% of the total amount of molecular hydrogen formed in the reaction zone A is formed. The method of claim 1 or 2, wherein the molecular hydrogen is burned together with the molecular oxygen in the reaction zone A when 65 mol% of the total amount of molecular hydrogen formed in the reaction zone A is formed. The method of claim 1 or 2, wherein the steam present in the product gas A in the separation zone I is at least partially removed and the removal is achieved by the product gas A intermingling and direct cooling. The method of claim 1 or 2, wherein a portion of the residual gas system having the residual gas composition is recycled to the reaction zone A as a feed stream comprising propane, and the residual gas comprises molecular oxygen. The method of claim 1 or 2, wherein the molecular oxygen sent to the reaction zone A is also fed as a component of air. The method of claim 1 or 2, wherein the molecular oxygen sent to the reaction zone A is also fed as a component comprising no more than 50% by volume of a component other than molecular oxygen. The method of claim 1 or 2, wherein the removal of the propane from the absorbent in the separation zone IV is carried out by stripping with a gas comprising steam and the stripping gas system supporting the propane as a feed comprising propane The stream is recycled to reaction zone A. The method of claim 1 or 2, wherein the reaction zone A is in an adiabatic configuration. The method of claim 1 or 2, wherein 10 to 40 mol% of the total amount of propane fed to the reaction zone A is converted in the reaction zone A by dehydrogenation in a single pass through the reaction zone A. The method of claim 1 or 2, wherein the reaction zone A comprises at least one disk reactor. The method of claim 1 or 2, wherein the reaction gas A formed in the reaction zone A has the following contents: 50 to 80% by volume of propane, 0.1 to 20% by volume of propylene, and 0 to 10% by volume of H ₂ , 0 to 20% by volume of N ₂ , and 5 to 15% by volume of H ₂ O. The method of claim 1 or 2, wherein the reaction gas A formed in the reaction zone A has the following contents: 55 to 80% by volume of propane, 0.1 to 20% by volume of propylene, and 2 to 10% by volume of H ₂ , 1 to 5 vol% of O ₂ , 0 to 20 vol% of N ₂ , and 5 to 15 vol% of H ₂ O. The method of claim 1 or 2, wherein the product gas A withdrawn from the reaction zone A has the following contents: 30 to 50% by volume of propane, 15 to 30% by volume of propylene, and 0 to 10% by volume of H ₂ 10 to 25% by volume of H ₂ O, 0 to 1% by volume of O ₂ , and 0 to 35% by volume of N ₂ . The method of claim 1 or 2, wherein the product gas A is cooled by indirect heat exchange with the feed stream comprising propane that has been removed in the separation zone IV and recycled to the reaction zone A. The method of claim 1 or 2, wherein the heterogeneously catalyzed partial gas phase oxidation It is a two-stage partial oxidation of propylene to acrylic acid. The method of claim 27, wherein the reaction gas sent to the second oxidation stage has the following contents: 3 to 25% by volume of acrolein, 5 to 65% by volume of molecular oxygen, and 6 to 70% by volume of propane. 0 to 20% by volume of H ₂ , 8 to 65% by volume of H ₂ O, and 0 to 70% by volume of N ₂ . The method of claim 1 or 2, wherein the residual gas has the following contents: 1 to 20% by volume of H ₂ O, 0 to 80% by volume of N ₂ , 10 to 90% by volume of propane, 0 to 20 by volume % H ₂ , 0 to 10% by volume of O ₂ , 1 to 20% by volume of CO ₂ , and 0 to 5% by volume of CO. The method of claim 1 or 2, wherein the feed stream comprising propane which has been removed in the separation zone IV and recycled to the reaction zone A has the following inclusions: 80 to 99.99 mol% of propane, 0 to 5 mol% of propylene, and 0 to 20 mol% of H ₂ O. The method of claim 1 or 2, wherein a part of the residual gas system having the residual gas composition is recycled to the reaction zone A as a feed stream comprising propane, and the partial residual gas is based on the total amount of the residual gas Up to 10% by weight. The method of claim 1 or 2, wherein the molecular oxygen fed to the reaction zone B is also fed as a component which does not include more than 50% by volume of a component other than molecular oxygen. The method of claim 1 or 2, wherein the molecular oxygen sent to the reaction zone B is also fed as a component of air.