Process for the continuous production of either acrolein or acrylic acid as target product from propene

Abstract

1. A process for the continuous production of either acrolein or acrylic acid as a target product from propene, comprising the steps of producing a product gas containing the target product by catalyzed gas-phase partial oxidation of propene, transferring the target product from the product gas to a liquid phase in a separation zone, and conducting from the separation zone a residual gas stream, a major portion of which is returned to the partial oxidation and the remainder of which stream is purged from the process as an off-gas, from which synthesis gas can be produced or which can otherwise be added to the produced synthesis gas.

Description

[Technical field]

[0001] The present invention relates to a process for the continuous production of either acrolein or acrylic acid as a target product from propene under steady-state operating conditions, comprising the steps of: A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 20% by volume of water, 0-20% by volume of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio of CO2:CO ranging from 1 to 4, and - Molecular oxygen and molecular nitrogen in a molar ratio of O2:N2 of at least 0.5 and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 30% by volume of water, 0-20% by volume of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio of CO2:CO ranging from 1 to 4, and - Molecular oxygen and molecular nitrogen in a molar ratio of O2:N2 of at least 0.25 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring to a separation zone either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid, in which the respective target products are separated from the respective product gas mixture by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone, which have a boiling point (herein the term boiling point also includes "or sublimation point") of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. The present invention relates to a method comprising the steps of: [Background technology]

[0002] In this document, continuous operation / production means that the initial reaction gas mixture 1 is continuously fed without interruption to the first reaction stage over a relatively long period of time, typically longer than one day, preferably at least several days (at least 2, or 3, or 4, or 5 days), and the target product is continuously removed from the separation zone in its desired purity.

[0003] Acrolein is an important intermediate for the preparation of, for example, glutaraldehyde, methionine, 1,3-propanediol, 3-picoline, folic acid and acrylic acid.

[0004] Acrylic acid is an important commodity chemical that is utilized as a monomer for preparing polymers that are used as binders, for example in dispersed distribution in aqueous media.Another field of use for acrylic acid polymers is superabsorbents in the hygiene sector and other application areas.

[0005] It is generally known that acrylic acid can be produced particularly selectively by heterogeneously catalyzed partial oxidation of propene in the gas phase using molecular oxygen over catalysts present in solid aggregated state. The active materials of the catalysts used are usually oxide compositions. Apart from oxygen, these usually contain more than one other element (multicomponent oxide compositions). Particularly frequently, the catalytically active oxide compositions used are those which contain more than one metal (especially transition metal) element. These are called multimetal oxide compositions. Multicomponent oxide compositions are usually not a simple physical mixture of the oxides of the elemental components, but a heterogeneous mixture of complex polycompounds of these elements.

[0006] As is generally known to those skilled in the art, the gas-phase partial oxidation of propene to acrylic acid proceeds essentially in two steps, sequentially along a reaction coordinate, which firstly leads to acrolein and secondly to acrylic acid from acrolein. As a result, if a method is suitable for the production of acrylic acid from propene by heterogeneously catalyzed gas-phase partial oxidation, it is also automatically suitable for the preparation of acrolein from propene, the production of acrylic acid being able to be stopped at the acrolein stage at any time.

[0007] Furthermore, the fact that the reaction proceeds in two chronologically successive steps makes it possible in a manner known per se to carry out the production of acrylic acid from propene by partial oxidation in the gas phase on heterogeneous catalysts in two successive oxidation stages, allowing optimal matching of the oxidation catalysts used in each of the two oxidation stages. Thus, for the first oxidation stage (propene → acrolein), catalysts based on multimetal oxides containing the element combination Mo-Bi-Fe are generally preferred, while for the second oxidation stage (acrolein → acrylic acid), catalysts based on multimetal oxides containing the element combination Mo-V are generally preferred. Suitable multimetal oxide catalysts for the two oxidation steps have been described many times and are well known to the person skilled in the art (e.g. EP-1005908, WO 00 / 53559, WO 00 / 53556, WO 00 / 53557, WO 02 / 24620, WO 2004 / 085368, WO 2004 / 085367, WO 2004 / 085365, WO 2004 / 085363, WO 2004 / 108284, WO 2004 / 108267, WO 2005 / 030393, WO 2005 / 047226, WO 2005 / 047224, WO 2005 / 049200, WO 2007 / 082827, WO 2008 / 087116, WO 2010 / 000764, WO 2010 / 028977, WO 2010 / 028991, WO 2010 / 066645, WO 2011 / 134932, WO 2012 / 049246, WO 2013 / 007736, WO 2013 / 167405, WO 2015 / 039982, WO 2015 / 067659, WO 2015 / 0128356 and DE-102018200841).

[0008] Usually, the active composition of multimetal oxide is not applied in powder form, but is formed into a certain catalyst shape. For example, unsupported catalysts can be prepared by compressing the powder of the active composition or its uncalcined and / or partially calcined precursor composition (in the latter two cases, the final calcination is performed after molding) into the desired catalyst shape (for example, tableting or extrusion), optionally with the addition of auxiliaries, for example graphite, hexagonal boron nitride or stearic acid as a lubricant and / or shaping aids and reinforcing agents, for example glass microfibers, asbestos, silicon carbide or potassium titanate. Examples of suitable unsupported catalyst shapes include solid or hollow cylinders and spheres.

[0009] The powdered active composition or its powdered precursor composition, uncalcined and / or partially calcined, can of course also be shaped by application onto a previously shaped inert catalyst support. The coating of the support body to produce the coated catalyst is generally carried out in a suitable rotatable vessel, as disclosed, for example, in DE-A 2909671, EP-A 293859 or EP-A 714700. To coat the support body, the powdered composition to be applied is appropriately moistened and, after application, is dried again, for example by means of hot air (for example, in the case of coating with an uncalcined powdered precursor composition, followed by calcination). Alternatively, the powdered active composition or its powdered precursor composition can also be advantageously coated onto a shaped support body (support) by a fluidized bed method, as described, for example, in WO 2005 / 030380. Useful support materials are those that behave substantially inertly with respect to both the respective target reaction underlying the production process and the side reactions accompanying the target reaction (see, for example, porous or non-porous aluminum oxide, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates, e.g., magnesium silicate or aluminum silicate). As an unsupported catalyst shape, the support body can have a regular or irregular shape, although regularly shaped support bodies with a well-defined surface roughness are preferred, e.g., spheres, solid or hollow cylinders.

[0010] In general, the heterogeneously catalyzed gas-phase partial oxidation of propene to acrolein and acrolein to acrylic acid is carried out at high reaction temperatures (usually several hundred degrees Celsius, usually between 250° C. and 500° C. for propene to acrolein and between 200° C. and 450° C. for acrolein to acrylic acid). The operating pressures (absolute) are usually in the range of 1 bar to 6 bar in both cases.

[0011] The relevant shaped catalyst body is usually part of a fixed catalyst bed through which the relevant reaction gas mixture containing the reactants molecular oxygen and propene or molecular oxygen and acrolein is passed at the relevant reaction temperature and operating pressure. In the fixed catalyst bed through which the reaction mixture passes, the target reaction takes place during the residence time of the reaction mixture. In the propene to acrolein step, the propene conversion in a single pass is usually between 80 mol % and 99 mol %, with an associated selectivity S for the formation of the target product acrolein. A is typically between 70 mol% and 99 mol%. In the acrolein to acrylic acid stage, the acrolein conversion in a single pass is typically between 95 mol% and 99.99 mol%, and the associated selectivity S for the formation of the target product acrylic acid, based on the converted propene, evaluated over both reaction stages, is AA is 80 mol % to 99.9 mol %.

[0012] Thus, the simplest embodiment of the associated reaction (oxidation) stage comprises a tube bundle reactor in which contact tubes (reactor tubes) are filled with an associated fixed catalyst bed. For the purpose of technically advantageously controlling the reaction temperature in the reactor tubes, a fluid (e.g. liquid) heat exchange medium (e.g. salt melt) is passed through the space around the contact tubes (see, for example, EP-A 700714 and EP-A 700983). As an alternative, thermoplate reactors can also be used in which the fixed catalyst bed is present as a planar arrangement between cooling plates (see, for example, US-A 2005 / 0042154 and US-A 2005 / 0020851).

[0013] Likewise, the two oxidation stages can be introduced in the form of an oxidation reactor system comprising two oxidation reactors in series and, optionally, if desired, a heat exchanger between the two reactors for cooling the reaction gas mixture leaving the first reactor. Using this arrangement, the other reaction conditions (e.g. reaction temperature) can be optimally matched in a simple manner to the respective oxidation stage. In principle, the aforementioned arrangement allows the transfer of the product gas mixture of the first oxidation stage without any intermediate treatment to the second oxidation stage. However, the molecular oxygen required for the second oxidation stage is advantageously fed here only to the second oxidation reactor. In an alternative embodiment, the heterogeneously catalyzed partial oxidation of propylene to acrylic acid in the gas phase can also be carried out in a single reactor. In this case, for example, both reaction steps can occur in one oxidation reactor packed with a catalyst packing that changes abruptly along the reaction coordinate.

[0014] What is common to all possible design variants is that due to the highly exothermic nature of the partial oxidation of propene or acrolein, the oxidation reactor is usually fed with an initial reaction gas mixture containing the reactants molecular oxygen and propene or molecular oxygen and acrolein diluted with a gas that is essentially inert under the conditions of the respective gas-phase partial oxidation. In this context, an inert diluent gas is one whose constituents remain unchanged to the extent of more than 95 mol %, preferably more than 98 mol %, under the conditions of the heterogeneously catalyzed gas-phase partial oxidation in a single pass through the respective reaction stage. Usually, the inert diluent gas constitutes the major proportion by volume of the respective initial reaction gas mixture (feed reaction gas mixture). The action of the inert diluent gas is mainly based on the simple dilution of the reactants and their ability to absorb the heat of reaction. Examples for usable inert diluent gases are hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene, water (H2O), molecular nitrogen (N2), carbon oxides (CO, CO2) and any mixture of some or all of the above mentioned gases. As pure compounds, inert diluent gases having a boiling point below 250 K at a working pressure of 1 bar are preferred.

[0015] With respect to the stoichiometry of the respective partial oxidation step to obtain the desired target compound, it is generally necessary to use molecular oxygen used as oxidant in the respective supplied reaction gas mixture in at least stoichiometric or superstoichiometric amounts (e.g. to reoxidize the active components of the multimetal oxide of the applied catalyst and to reduce the carbon deposited on the surface of the catalyst).

[0016] The separation of the respective target product (acrolein or acrylic acid) from the respective product gas mixture resulting from the respective heterogeneously catalyzed partial gas phase oxidation is usually carried out in a corresponding separation zone by transferring at least 95 mol % of the respective target product contained in the respective product gas mixture from the gas phase to the liquid phase, which can be carried out as known from the prior art processes.

[0017] Typically, each product gas mixture entering the separation zone is first subjected to indirect and / or direct cooling (such cooling can already bring about the separation of constituents of the product gas mixture containing the target product, which are continuously transferred to the separation zone and which, as pure compounds, have a boiling point at an operating pressure of 1 bar higher than that of the target product) from said product gas mixture. The transfer of the target products present in the product gas mixture cooled in this way or present in the product gas mixture not cooled to the liquid phase can be carried out by absorption and / or condensation in a liquid absorbent (e.g. fractional condensation), for example as described in US-A 3 433 840, WO 2008 / 090190, WO 2011 / 000808, US-B 7183428, WO 2014 / 122044, US-A 5 315037, WO 2003 / 068721, DE-A 10213998 and DE-A 3 429 391. Useful absorbents for substantially both target products are either water or aqueous solutions or high-boiling organic solvents. Subsequent work-up of the resulting desired target product containing liquid phase in the separation zone via application of thermal separation processes such as stripping, distillation, rectification (including azeotropic rectification, respectively), extraction and / or crystallization ultimately produces the target product in its desired purity, which target product is usually continuously withdrawn from the separation zone to which the product mixture is also continuously fed.

[0018] Those constituents of the product gas mixture containing the target product, which are continuously transferred to the separation zone, which, as pure compounds, have a boiling point below the boiling point of pure water at an operating pressure of 1 bar, remain mainly in the gas phase in the separation zone (in the case of water itself and in the case of constituents which have a boiling point higher than the boiling point of pure water under corresponding conditions, the specific operating conditions applied to transfer the target product from the respective product gas mixture to the liquid phase determine the respective amounts of these constituents and water remaining in the residual gas mixture), and the total mass flow is contained in the residual gas mixture which is continuously led off from the separation zone. In particular, such compounds contained in the product gas mixture which is continuously transferred to the separation zone, which, as pure compounds, have a boiling point below 250 K at an operating pressure of 1 bar (such compounds include, above all, the respective components of the applied inert diluent gas), are the major part of the residual gas mixture. Typically, at least 95 mol-% of the total molar amount of those compounds which are components of the product gas mixture continuously transferred to the separation zone and which, as pure compounds, have a boiling point below 250 K at an operating pressure of 1 bar are ultimately contained in the total mass stream of the residue gas mixture continuously led off from the separation zone (under stationary operating conditions).

[0019] Usually, the residual gas mixture further contains residual unconverted propene. Since propene is a valuable raw material for the production process, and the above explanations naturally lead to an interest in recycling the entire mass flow rate of the residual gas mixture continuously led off from the separation zone to the propene-to-acrolein reaction stage as a feed charge for preparing the initial reaction gas mixture fed to the reaction stage. This helps, on the one hand, to increase both the overall ("in multiple passes") propene conversion and the yield of the target product (e.g. acrylic acid) resulting from the production process, and, on the other hand, provides a suitable source of inert diluent gas that can be used for preparing the initial reaction gas mixture.

[0020] However, it is not possible to completely recycle the relevant total mass flow rate of the residual gas mixture continuously withdrawn from the separation zone as a feed to the relevant starting reaction gas mixture, because, under stationary operating conditions, at least a balancing part of the total mass flow rate of the residual gas mixture continuously withdrawn from the separation zone must be continuously discharged from the production process as off-gas in order to prevent the accumulation of certain components of the residual gas mixture of the production process, for example continuously by-produced in a reaction stage (for example CO, CO2 or H2O) or otherwise continuously introduced into the production process (for example molecular nitrogen (N2) optionally contained as a companion to molecular oxygen (O2) in a source of fresh molecular oxygen (for example air) used for the preparation of the respective starting gas mixture continuously fed to the respective reaction stage). Technically advantageously, the off-gas has the same composition as the residual gas mixture, and only the difference between the total mass flow rate of the residual gas mixture continuously led off from the separation zone and the mass flow rate of the off-gas stream continuously discharged from the production process is recycled as a feed stream (returned as annular gas stream), which feed stream also consequently has the same composition as the residual gas mixture for preparing the associated starting reaction gas mixture stream.

[0021] For most of the processes recommended according to the prior art for producing either acrolein or acrylic acid as target product by heterogeneously catalyzed partial oxidation of propene in the gas phase, the feed gas stream used for preparing the stream of the initial reaction gas mixture fed to the propene to acrolein reaction stage consists of a stream of ambient air (based on its total amount containing about 78% by volume of N2 and about 21% by volume of O2), a stream of chemical grade propene (consisting of about 95% by volume of propene and about 5% by volume of propane) and a stream of the recycled part of the residual gas mixture resulting from this process. As a result, the mass flow rate of the off-gas stream in these processes is typically 38% or more of the total mass flow rate of the residual gas mixture continuously led off from the separation zone (see, for example, WO 2008 / 090190 or WO 2011 / 000808). In addition, the molecular nitrogen (N2) content of the off-gas in the aforementioned production methods is usually more than 85% by volume, and such off-gas is not further used as a raw material for another synthesis, mainly for economic reasons, but is usually only disposed of via combustion in order to at least benefit from the resulting combustion energy (see WO 97 / 48669). The disadvantage of this procedure is that the combustion of the carbon-containing constituents of the off-gas produces additional carbon dioxide in addition to the carbon dioxide already contained in the off-gas before its combustion, and the emission of carbon dioxide to the outside air worsens the carbon dioxide coverage of the process for producing either acrolein or acrylic acid as the target product from propene.

[0022] EP-A 731080 is directed to a process for the continuous heterogeneously catalyzed partial gas-phase oxidation of propene in an oxidation reactor to produce, inter alia, either acrolein or acrylic acid as the target product. In the EP-A 731080 production process, the starting reaction gas mixture fed to the oxidation reactor, apart from propene and molecular oxygen, preferably consists of more than 85% by volume, and in the best case 100% by volume, of an inert diluent gas consisting of combustible saturated hydrocarbons (except in the case of the processes of WO 2008 / 090190 or WO 2011 / 000808, the molecular oxygen required for such EP-A 731080 processes is taken from an essentially pure molecular oxygen source). According to the teaching of EP-A 731 080, after a single pass of the reaction gas mixture through the oxidation reactor, the aforementioned combustible constituents of the inert diluent gas present in the product gas stream leaving the oxidation reactor are not recycled to the heterogeneous gas-phase partial oxidation at all, but are used further (preferably as is) for the production of synthesis gas (one disadvantage of such a procedure is that the total amount of saturated hydrocarbons introduced into the starting reaction gas mixture via a separate feed stream is used only in a single pass as inert diluent gas). Thus, if the residual gas mixture continuously withdrawn from the separation zone is fed for such further use, this improves the carbon dioxide coverage of the process of EP-A 731080 over that of the processes of WO 2011 / 000808, respectively WO 2008 / 090190, but in such a case the unconverted propene still contained in said residual gas mixture is also not recycled to the oxidation reactor, which disadvantageously reduces both the overall propene conversion and the resulting yield of the target product of the production process (e.g. acrylic acid).If instead the unreacted propene still contained in said residual gas mixture is separated before the relevant further use and recycled separately to the oxidation reactor in order to increase the overall conversion of propene, this requires additional efforts and significantly disadvantageously reduces the efficiency of the production process of EP-A 731080.The further hint given in EP-A 731080 that optionally non-combustible inert diluent gas constituents (including those only formed as by-products in the course of the reaction), e.g. N2, CO2 and H2O, can also be separated and released to the atmosphere before further relevant use, is neither efficient nor beneficial with regard to the target of improving the carbon dioxide coverage of the method.

[0023] US-A 5,684,188 is directed to a process for the continuous heterogeneously catalyzed partial oxidation of propene to acrolein, acrylic acid or mixtures thereof in an oxidation reactor, the starting reaction gas (feed gas) mixture of which, apart from propene and molecular oxygen as oxidant, comprises only at least one essentially inert diluent gas under heterogeneously catalyzed gas phase oxidation conditions, in which, in continuous operation, at least a portion of the essentially inert diluent gas constituents present in the product gas mixture are separated therefrom and recycled as constituents of the feed to the oxidation reactor, in which, in continuous operation, the essentially inert diluent gas mixture of the feed gas mixture comprises more than 85% by volume, preferably at least 99% by volume, and in the best case 100% by volume, of at least one saturated hydrocarbon having 1 to 5 carbon atoms.

[0024] However, a very particular aspect of the teaching of US-A 5,684,188 is that the saturated hydrocarbons used as inert diluent gas constituents of the initial reaction gas mixture are considered valuable materials, unlike inert diluent gas constituents such as N2, H2O and CO2. Thus, according to US-A 5,684,188, a prerequisite for the economical operation of the process is to recycle the saturated hydrocarbons present in the product gas mixture as feed for the initial reaction gas mixture. Consequently, from the point of view of US-A 5,684,188, for the purpose of this recycle, the saturated hydrocarbons must be separated not only from the target products in the product gas mixture, but also from by-product gases such as CO2 and H2O. Furthermore, according to US-A 5,684,188, unreacted starting materials (propene) and / or intermediates (acrolein) still present in the product gas mixture can also be separated and recycled to the oxidation reactor as feed for the preparation of the initial reaction gas mixture. However, such separation of the hydrocarbons from the other constituents of the product gas mixture disadvantageously reduces the economic efficiency of the process. [Prior art documents] [Patent documents]

[0025] [Patent Document 1] EP-1005908 [Patent Document 2] WO 00 / 53559 [Patent Document 3] WO 00 / 53556 [Patent Document 4] WO 00 / 53557 [Patent Document 5] WO 02 / 24620 [Patent Document 6] WO 2004 / 085368 [Patent Document 7] WO 2004 / 085367 [Patent Document 8] WO 2004 / 085365 [Patent Document 9] WO 2004 / 085363 [Patent Document 10] WO 2004 / 108284 [Patent Document 11] WO 2004 / 108267 [Patent Document 12] WO 2005 / 030393 [Patent Document 13] WO 2005 / 047226 [Patent Document 14] WO 2005 / 047224 [Patent Document 15] WO 2005 / 049200 [Patent Document 16] WO 2007 / 082827 [Patent Document 17] WO 2008 / 087116 [Patent Document 18] WO 2010 / 000764 [Patent Document 19] WO 2010 / 028977 [Patent Document 20] WO 2010 / 028991 [Patent Document 21] WO 2010 / 066645 [Patent Document 22] WO 2011 / 134932 [Patent Document 23] WO 2012 / 049246 [Patent Document 24] WO 2013 / 007736 [Patent Document 25] WO 2013 / 167405 [Patent Document 26] WO 2015 / 039982 [Patent Document 27] WO 2015 / 067659 [Patent Document 28] WO 2015 / 0128356 [Patent Document 29] DE-102018200841 [Patent Document 30] DE-A2909671

Patent document 31

Patent document 32

Patent document 33

Patent document 34

Patent document 35

Patent document 36

Patent document 37

Patent document 38

Patent Document 39

Patent document 40

Patent document 41

Patent document 42

Patent document 43

Patent document 44

Patent document 45

Patent document 46

Patent document 47

Patent document 48

Patent document 49

Patent document 50

Patent document 51

Patent document 52

Patent document 53

Patent document 54

Patent document 55

Patent document 56

Patent document 57

Patent document 58

Patent document 59

Patent document 60

Patent document 61

Patent document 62

Patent document 63

Patent document 64

Patent document 65

Patent document 66

Patent document 67

Patent document 68

Patent document 69

Patent document 70

Patent document 71

Patent document 72

Patent document 73

Patent document 74

Patent document 75

Patent document 76

Patent document 77

Patent document 78

Patent document 79

Patent document 80

Patent document 81

Patent document 82

Patent document 83

Patent document 84

Patent document 85

Patent document 86

Patent document 87

Patent document 88

Patent document 89

Patent document 90

Patent document 91

Patent document 92

Patent document 93

Patent document 94

Patent document 95

Patent document 96

Patent document 97

Patent document 98

Patent document 99

Patent document 100

Patent Document 101

Patent document 102

Patent document 103

Patent document 104

Patent document 105

Patent document 106

Patent document 107

Patent document 108

Patent document 109

Patent document 110

Patent document 111

Patent document 112

Patent document 113

Patent document 114

Patent document 115

Patent document 116

Patent document 117

Patent document 118

Patent document 119

Patent document 120

Patent document 121

Patent document 122

Patent document 123

Patent document 124

Patent document 125

Patent document 126

Patent document 127

Patent document 128

Patent document 129

Patent document 130

Patent document 131

Patent document 132

Patent document 133

Patent document 134

Patent document 135

Patent document 136

Patent document 137

Patent document 138

Patent document 139

Patent document 140

Patent document 141

Patent document 142

Patent document 143

Patent document 144

Patent document 145

Patent document 146

Patent document 147

Patent document 148

Patent document 149

Patent document 150

Patent document 151

Patent document 152

Patent document 153

Patent document 154

Patent document 155

Patent document 156

Patent document 157

Patent document 158

[0026] [Non-Patent Document 1] Chem. Technik, 50. Jg. Heft 5, September / Oktober1998, pp. 241-245 [Non-Patent Document 2] Ullmann's Encyclopedia of Industrial Chemistry, Fifth Completely Revised Edition, Volume A12 (Formamides to Hexamethylendiamine), VCH Verlagsgesellschaft mbH, D-6940 Weinheim, pages 178 and 179, section 1.3. [Non-Patent Document 3] Ullmann's Encyclopedia of Industrial Chemistry, Fifth Completely Revised Edition, Volume A12 (Formamides to Hexamethylendiamine), VCH Verlagsgesellschaft mbH, D-6940 Weinheim, pp. 192-194, Section 2.2.3. Summary of the Invention [Problem to be solved by the invention]

[0027] It is therefore an object of the present invention to provide a process for the continuous production of either acrolein or acrylic acid as target product by heterogeneously catalyzed partial oxidation of propene in the gas phase and subsequent separation of the target product from the respective product gas mixture in a separation zone by transferring the target product from the product gas mixture to the liquid phase, the off-gas having a production and composition such as to allow its further use (preferably as is) as a feedstock for the production of synthesis gas (also called "syngas") and / or for addition to already produced synthesis gas, if desired, improving the carbon dioxide balance of the process, without causing / without significant economic disadvantages with respect to its production and / or use, other than in the processes of US-A 5684188, WO 2008 / 090190, WO 2011 / 000808 and EP-A 731080.

[0028] Another object of the present invention is to provide a process which further allows for both high overall propene conversion and resulting high yield of the target product, such as acrylic acid.

[0029] It is a further object of the present invention to provide a method which, in a preferred embodiment, is carried out outside the explosion zone. [Means for solving the problem]

[0030] The inventors have discovered that all these objects are achieved by a process for the continuous production of either acrolein or acrylic acid as target product from propene under steady-state operating conditions, comprising: A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 20% by volume of water, 0 to 20% by volume (preferably 0 to 10%) of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid; It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, - and molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5 and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 30% by volume of water, 0 to 20% by volume (preferably 0 to 10%) of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio of CO2:CO ranging from 1 to 4, and - Molecular oxygen and molecular nitrogen in a molar ratio of O2:N2 of at least 0.25 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid into a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target product contained in the respective product gas mixture from the gas phase into the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % (preferably at least 96 mol %, particularly preferably at least 97 mol %, in the best case at least 98 mol %, or at least 99 mol %, or at least 99.5 mol %, but usually less than 100 mol %) of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone which have a boiling point below 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. Including, a) the mass flow rate of the recycle gas stream P is at least 65% but not more than 99.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 35% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.01 to 1% by volume of propene, 1 to 10% by volume of molecular oxygen, 0.5 to 25% by volume of water, 0 to 20% by volume (preferably 0 to 10%) of molecular nitrogen, 0.5 to 40% by volume of carbon monoxide, 1 to 75% by volume of carbon dioxide, 0-92% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane, and ethene; and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid It has been found that this can be achieved by a method in which

[0031] According to the invention, advantageous reaction temperatures in the first reaction stage are between 280°C and 480°C, preferably between 290°C and 470°C, more preferably between 300°C and 450°C. The operating pressure in the first reaction stage of the process of the invention is preferably between 1 and 5 bar, more preferably between 1 and 4 bar, most preferably between 1 and 3.5 bar. The propene conversion in a single pass in the first reaction stage of the process of the invention is preferably in the range of 85 mol% to 99 mol%. Better said conversion is between 90 mol% and 99 mol%, best said conversion is between 92 mol% and 98.5 mol%. The associated selectivity S of the formation of the target product acrolein in the first reaction stage is preferably between 1 and 5 bar, more preferably between 1 and 4 bar, most preferably between 1 and 3.5 bar. A is advantageously between 75 mol % and 99 mol %, better still between 80 mol % and 99 mol %.

[0032] According to the invention, advantageous reaction temperatures in the second reaction stage are between 210°C and 430°C, preferably between 220°C and 410°C, more preferably between 230°C and 390°C. The operating pressure in the second reaction stage of the process of the invention is preferably between 1 and 5 bar, more preferably between 1 and 4 bar, most preferably between 1 and 3.5 bar. The acrolein conversion in a single pass in the second reaction stage of the process of the invention is preferably between 97 mol% and 99.95 mol%. Better said conversion is between 98 mol% and 99.9 mol%, best said conversion is between 99 mol% and 99.8 mol%. The associated selectivity S of the formation of the target product acrylic acid, based on the converted propene, evaluated over both reaction stages, is AAis advantageously between 83 mol % and 99 mol %, better between 85 mol % and 99 mol %, and best between 87 mol % and 99 mol %.

[0033] Taking into account the envisaged subsequent optional use of the off-gas of the inventive process for the production of synthesis gas, ambient air, due to its high molecular nitrogen content, is not suitable as a source of molecular oxygen required as oxidant for the purposes of the inventive process (e.g. preparing (forming, producing) the starting reaction gas mixture 1) (including the optional step of adding a molecular oxygen-containing secondary gas to the product gas mixture 1 for preparing the starting reaction gas mixture 2 between the first and second reaction stages of the inventive process). Thus, in contrast, a mixture of molecular oxygen and an inert gas containing at least 90% by volume, preferably at least 95% by volume, more preferably at least 98% by volume, even better at least 99% by volume, best between 99.4% by volume and 99.9% by volume or between 99.5% by volume and 99.8% by volume of molecular oxygen is a suitable oxygen source for the aforementioned purposes. Air with molecular nitrogen removed is such a mixture (compared to using ambient air as the oxygen source, the use of molecular oxygen with a purity of at least 90% by volume O2 and 10% by volume or less N2 also has the advantage that the quality of the oxygen source is decoupled and no longer influenced by unfavorable environmental influences at the respective production site (e.g. no corrosive damage due to the chloride content in the ambient air, plus the compressor required to compress the ambient air to the working pressure can be omitted)). Thus, (molecular) oxygen with purities of at least 90% by volume O2 and not more than 10% by volume N2, or at least 95% by volume O2 and not more than 5% by volume N2, or at least 98% by volume O2 and not more than 2% by volume N2, or at least 99% by volume O2 and not more than 1% by volume N2, or 99.4% to 99.9% by volume O2 and 0.1% to 0.6% by volume N2, or 99.5% to 99.8% by volume O2 and 0.2% to 0.5% by volume N2 are thus suitable as oxygen sources.

[0034] Of course, pure molecular oxygen (100% O2 by volume purity) can also be used as an advantageous oxygen source for the process of the present invention.

[0035] As possible sources of starting material propene used to prepare the initial reaction gas mixture 1 (propene source, raw propene, propene feedstock), polymer grade propene (containing ≧99% by volume of propene and ≦1% by volume of propane), chemical grade propene (containing ≧90% by volume of propene and ≦10% by volume but ≧2% by volume of propane) and refinery grade propene (containing ≧70% by volume of propene and ≦30% by volume but ≧15% by volume of propane) can be applied (for example as specifically defined in WO 2004 / 007405). Since a part of the propane accompanying the propene in the aforementioned propene source also appears eventually in the off-gas of the process of the present invention, and since the heterogeneously catalyzed steam reforming of saturated hydrocarbons is a suitable means of producing synthesis gas, chemical grade propene is a very preferred propene source for the process of the present invention.

[0036] Preferably, in the process of the invention, the mass flow rate of the recycle gas stream P is at least 70% but not more than 99.5% of the mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone, and the mass flow rate of the off-gas stream O is not more than 30% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone.

[0037] Particularly preferred in the process of the invention is a mass flow rate of the recycle gas stream P which is at least 80% but not more than 99.0% of the mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone, and a mass flow rate of the off-gas stream O which is not more than 20% but at least 1.0% of the total mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone.

[0038] Even more preferably in the process of the invention the mass flow rate of the recycle gas stream P is at least 90% but not more than 99.0% of the mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone, and the mass flow rate of the off-gas stream O is not more than 10% but at least 1.0% of the total mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone.

[0039] In the process of the invention, the most preferred mass flow rate of the recycle gas stream P is at least 95% but not more than 98.0% of the mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone, and the mass flow rate of the off-gas stream O is not more than 5% but at least 2.0% of the total mass flow rate M of the residual gas mixture R continuously withdrawn from the separation zone.

[0040] A first highly preferred embodiment of the process of the present invention is a process for the continuous production of either acrolein or acrylic acid as target product from propene under stationary operating conditions. A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 15 to 25% by volume of carbon monoxide, 20 to 55% by volume of carbon dioxide, 10-30% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially propane only); and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, and molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5. and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 20% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 15 to 25% by volume of carbon monoxide, 20 to 55% by volume of carbon dioxide, 10-30% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially propane only); and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio of CO2:CO ranging from 1 to 4, and - Molecular oxygen and molecular nitrogen in a molar ratio of O2:N2 of at least 0.25 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. Including, a) the mass flow rate of the recycle gas stream P is at least 90% but not more than 99% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 10% but at least 1% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.05 to 0.5% by volume of propene, 1-4% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 15 to 30% by volume of carbon monoxide, 30-60% by volume of carbon dioxide, 10-50% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially only propane); and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid That is, the method.

[0041] In the case of the first highly preferred embodiment mentioned above, the feed streams forming (used to prepare) the initial reaction gas mixture 1 preferably consist exclusively of a feed stream of chemical grade propene (propene feedstock containing ≧90 vol.% propene and ≦10 vol.% but ≧2 vol.% propane), a stream of molecular oxygen with a purity of ≧99.0 vol.% and ≦99.9 vol.% O2 and ≦1 vol.% and ≧0.1 vol.% N2, and a recycle gas stream P.

[0042] A second highly preferred embodiment of the process of the present invention is a process for the continuous production of either acrolein or acrylic acid as target product from propene under steady state operating conditions, comprising: A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 1.2 to 15 volume percent carbon monoxide, 3 to 30% by volume of carbon dioxide, 40-80% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol % of the total molar amount of these hydrocarbons is methane), and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, - molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5, and - Carbon monoxide and total molar amount of hydrocarbons in the CO:total hydrocarbons molar ratio range of >0-0.3 and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 20% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 1.2 to 15 volume percent carbon monoxide, 3 to 30% by volume of carbon dioxide, 40 to 80% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol % of the total molar amount of these hydrocarbons is methane), and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, - molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.25, and - Carbon monoxide and total molar amounts of hydrocarbons (including acrolein) in the CO:total hydrocarbons molar ratio range of >0-0.3 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing a gas stream P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas stream P from the production process as an off-gas stream O having the same composition as the residual gas mixture R, Including, a) the mass flow rate of the recycle gas stream P is at least 75% but not more than 98% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 25% but at least 2% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.05 to 0.8 volume percent propene, 1-4% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 1.2 to 15% by volume of carbon monoxide, 3 to 30% by volume of carbon dioxide, 45-90% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol % of the total molar amount of these hydrocarbons is methane), and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid That is, the method.

[0043] In the case of the second highly preferred embodiment mentioned above, the feed streams forming (used to prepare) the initial reaction gas mixture 1 preferably consist exclusively of a propene feed stream of polymer grade propene (propene feedstock containing ≧99 vol.% propene and ≦1 vol.% propane) or chemical grade propene (propene feedstock containing ≧90 vol.% propene and ≦10 vol.% but ≧2 vol.% propane), a stream of molecular oxygen with a purity of ≧99.0 vol.-% and ≦99.9 vol.-% O2 and ≦1 vol.-% and ≧0.1 vol.-% N2, a recycle gas stream P, and an additional stream of at least one hydrocarbon selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol.% (better at least 90 mol.%, more preferably at least 95 mol.%, best case 100 mol.%) of the total molar amount of these hydrocarbons being methane or natural gas).

[0044] In principle, the stationary production method according to the invention can be carried out with non-pyrophoric (non-explosive) and pyrophoric (explosive) gas mixtures (for the ignition properties of gas mixtures, see US-A 6090977 and US-A 5684188), provided that in both cases sufficient safety measures are taken (see WO 2004 / 007405, page 7, lines 21 to 31). What is important to answer the question whether a gas mixture is explosive (pyrophoric) is whether a combustion (ignition, explosion) initiated by a local ignition source (e.g. a burning platinum wire) spreads in the gas mixture present under certain conditions (pressure, temperature) (see DIN 51649 and WO 04 / 007405, in particular the experimental description on page 2). If there is a spread, the mixture is classified here as explosive (pyrophoric). If there is no spread, the mixture is classified here as non-explosive (non-pyrophoric). As disclosed in WO 04 / 007404, it is easily possible to determine the explosion limit line for some compounds that may constitute various gas mixtures exhibiting different compositions at a certain temperature and pressure. This is the line that separates explosive gas mixtures from non-explosive gas mixtures in an explosion diagram (showing explosion behavior vs. composition). In this document, it is considered that if the initial reaction gas mixture 1 is generally non-explosive under the relevant stationary operating conditions (which usually contains the highest concentration of molecular oxygen; sections

[0028] ,

[0029] and

[0030] of US-A 2004 / 0015012 apply accordingly), this also generally applies to the reaction gas mixture formed (produced) within the process of the invention. Preferably, the process of the invention is carried out outside the explosive (ignitable) range. From an economic point of view, it is preferred to carry out the process of the invention by applying an initial reaction gas mixture 1 which is non-explosive under reaction conditions on the one hand and has a high concentration of the reactants propene and molecular oxygen on the other hand (this applies correspondingly to the reactants acrolein and molecular oxygen in the initial reaction gas mixture 2). All the above-mentioned ratios Q=(n CO ) / (n CO +n 炭化水素) is preferably in the range of >0 to 0.8, more preferably 0.005 to 0.7, and best in the range of 0.005 to 0.6 (n CO means the total molar amount of carbon monoxide contained in the initial reaction gas mixture 1, n 炭化水素 Q = (n CO ) / (n CO +n 炭化水素 )>0-0.8 (more preferably 0.005-0.7, and best 0.005-0.6) also applies to the starting reaction gas mixture 2. However, for the starting reaction gas mixture 2, n 炭化水素means the total molar amount of hydrocarbons (saturated and unsaturated) plus the total molar amount of acrolein contained in the initial reaction gas mixture 2. The above formula first takes into account that gases such as CO2, N2, H2O and all noble gases, under normal conditions, are gases in which their mixture with molecular oxygen (O2) in any mixture ratio is not capable of diffusing combustion reactions from an ignition source (hence such gases are designated as "non-flammable gases"), whereas carbon monoxide (CO), in its mixture with O2, is a flammable gas in a certain mixture ratio capable of diffusing combustion reactions from an ignition source. In addition, this equation takes into account that the desired and catalytically promoted target oxidation of propene to acrolein, respectively acrolein to acrylic acid (first reaction) of the process of the invention is always accompanied by undesired secondary reactions, the main of which is the combustion reaction of propene and acrolein to form carbon monoxide and carbon dioxide, and that, based on the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, in the process of the invention, compared to the state of the art, an increase in the proportion of the recycle gas stream P correspondingly increases the concentration of carbon monoxide in the starting reaction gas mixture 1 and starting reaction gas mixture 2, promoting the risk of deflagration. Finally, the above equation also takes into account the well-known fact from the prior art that such flammable behavior can be limited via the presence of fuel gases, for example saturated and unsaturated hydrocarbons in the starting reaction gas mixture 1, respectively 2 (see Chem. Technik, 50. Jg. Heft 5, September / October 1998, pp. 241-245).

[0045] What is important in the case of continuous operation of the method according to the invention with a composition of the starting reaction gas mixture 1 outside the explosive zone (since in this case the safety measures usually applied in other respects are relatively limited) is that the continuous operation has a cut-off mechanism which prevents the composition of the starting reaction gas mixture 1 outside the explosive zone from passing into the explosive zone during continuous operation (especially in the steady state) as a result of an accidental lapse in the supply of gas flows to form the starting reaction gas mixture 1. This cut-off respective shut-down mechanism can be designed in the following simple manner according to the teachings of WO 2004 / 007405, WO 2015 / 104397 and WO 2018 / 029215: a) storing in a computer an explosion diagram (the pressure and temperature dependence of the explosion diagram is limited and points where the changes are not too large can be ignored; section

[0122] of US 2004 / 0015012 applies accordingly) characterizing the initial reaction gas mixture 1 (as long as the specific heats of the non-flammable inert diluent gases do not differ significantly from each other, they affect the explosion diagram in the same way) in which explosive and non-explosive compositions are plotted against each other as a function of the composition of the initial reaction gas mixture 1 (for safety reasons, it may be appropriate to enter into the computer as the explosion diagram not the curve of the experimentally determined explosion limit, but rather a switching line which, in comparison with this, is shifted downwards by a safety margin (see section

[0065] of US-A 2004 / 0015012); b) the data set is continuously determined by determining the amount and optionally the composition of the gas feed streams forming the starting reactant gas mixture 1; c) the computer continuously calculates the instantaneous action points of the initial reactant gas mixture 1 on the explosion diagram from the data set obtained in b); d) If the distance of the application point from the nearest explosion limit is less than a specified minimum value, the flow of the gas stream supplying the initial reaction gas mixture 1 is automatically interrupted (closed off).

[0046] Preferably, the feed flows are not shut off at the same time. Rather, the order of shutting off the gas feed flows is preferably selected as follows (for safety reasons): 1. Feed stream containing fresh molecular oxygen = feedstock (source) of molecular oxygen (including any optional addition of a secondary gas containing molecular oxygen to the product gas mixture 1 between the first reaction stage and the second reaction stage to form the starting reaction gas mixture 2); 2. Fresh propene containing feed stream = propene feedstock; 3. Another hydrocarbon feed stream, which is optionally co-used as an inert diluent gas to the starting reactant gas mixture 1; 4. Optionally, a feed stream of water vapor, which can also function as an inert diluent gas in the initial reactant gas mixture 1 and is further known as a means for increasing the catalytic activity of the multimetal oxide catalyst used in the production process, is used simultaneously; and 5. Recycle gas stream P.

[0047] Such a shut-down mechanism is also suitably incorporated with respect to the initial reactant gas mixture 2.

[0048] In addition, CO and n 炭化水素 can be continuously determined for a given starting reaction gas mixture 1 and a given starting reaction gas mixture 2, the obtained values ​​being continuously fed to a computer, which calculates the ratio Q=(n CO ) / (n CO +n 炭化水素 ) is continuously calculated. If either Q approaches the value 0.7 or 0.8, the flow of the gas streams (e.g., any secondary molecular oxygen-containing gas streams optionally added between the first and second reaction stages) supplied to starting reaction gas mixture 1, respectively starting reaction gas mixture 2, is automatically disconnected (shut off) in the same manner as described above.

[0049] To generate (form) the stream of the initial reaction gas mixture 1, the individual feed streams are usually fed to the first reaction stage (first oxidation reactor) so that they meet from separate feed lines (usually all feed streams are already essentially compressed at the inlet pressure required to enter the first reaction stage) and are first fed to a mixer (for example a static mixer (see chamber with internal structures that generate turbulence), where the feed streams are mixed homogeneously and then, if necessary, after heating to the inlet temperature, are fed to the first reaction stage (first oxidation reactor). The entry of the individual feed gases into the lines leading to the static mixer (flow direction along said lines) is preferably performed in the opposite order (reversed order) (i.e. first (1.) the recycle gas stream P and last (5.) the feed stream containing fresh molecular oxygen) as the cut-off of the gas feed streams in case of cut-off respectively shutdown of the production process, for safety reasons.

[0050] However, even if the selected composition of the homogeneously mixed initial reaction gas mixture 1 is non-explosive, the problem remains: at the (local) point where a fresh molecular oxygen-containing feed stream is added (at such point, even if such a feed stream is added to the (feed) stream Z containing the recycle gas stream P, the optionally simultaneous feed stream of steam, the optionally simultaneous feed stream of any of the simultaneously used inert hydrocarbons other than propene, and the fresh propene feed stream (propene feedstock) are already perfectly homogeneously mixed), the locally generated gas mixture is usually explosive (this problem exists even if the outside air is used as the source of fresh oxygen, but due to the relatively high content of molecular nitrogen in air, this problem is less severe (less important)).

[0051] Therefore, according to the invention, the following procedure is advantageously used: First, a homogeneously mixed feed stream Z is generated using a simple static mixer, as described above. Then, with the help of a mixer that is particularly suited with respect to the problem to be solved, i.e. optimally protecting against the local formation of explosive gas mixtures, this stream is mixed with a feed stream (source) containing fresh molecular oxygen (a stream of ≧90% by volume O2 (preferably ≧95% by volume O2) and ≦10% by volume N2 (preferably ≦5% by volume N2)). After leaving the mixer, the resulting initial reaction gas mixture 1 is first fed to a reaction stage (respective oxidation reactor).

[0052] Such devices suitable for the relevant purposes are disclosed in WO 2009 / 078898 and US-B8334395 respectively. When using this device for the relevant mixing, a gas feed stream Z is mixed with a fresh molecular oxygen feed stream (having a source of molecular oxygen) in the presence of a coarse water droplet environment in a gas mixer. The water droplets surround and contact entrained particles in either the fresh molecular oxygen containing gas stream or the feed stream Z. The water acts to suppress, prevent and quench ignition of hydrocarbon and carbon monoxide gases in the mixer which would otherwise be caused by energetic collisions between such particles and structures in the gas mixer.

[0053] WO 2007 / 071737 discloses an alternative mixing device suitable for the relevant purpose. Performing the mixing in this mixing device involves conveying a gas feed stream Z as a first stream and a gas stream containing fresh molecular oxygen (source of molecular oxygen) as a second stream to a microchannel device, whereby mixing occurs and the mixture is drawn off. The advantage of this mixing device is that it allows for a mixing process that minimizes the length of time that the mixture containing molecular oxygen / (hydrocarbon+CO) can be present in the relevant explosive region. The gases are mixed at the "micro-level", i.e. on a very small scale, within the microchannels of the process of the microchannel device. After mixing of both feeds, pockets of oxygen-rich and oxygen-poor mixtures will naturally exist initially, but through the division and recombination of the oxygen stream within the microchannels of the process and the microchannel openings, if present, an average oxygen concentration below the explosive limit is established. As the gas mixture progresses through the microchannel device, these pockets disappear and the gases are well mixed at the micro-level.

[0054] Further suitable mixing devices are disclosed in US-B8404189 and WO 2009 / 078899 respectively. This gas mixer comprises a hollow pipe located inside a vessel containing a feed stream Z (first gas stream) and a vessel containing a feed stream of fresh molecular oxygen (second gas stream). The internal pipe further comprises a mixer tip at its distal end. The mixer tip comprises a body having an internal passage for directing the second gas from the pipe and an opening for introducing the second gas stream into the first gas stream in a radial plane at an acute angle to the longitudinal axis of the pipe. The pipe further comprises a deflection device on its outer surface longitudinally aligned with the opening of the mixer tip. The deflection device serves to deflect any particles entrained in the first gas stream away from the mixing zone where the two streams mix, minimizing the risk of ignition of the gas Z containing hydrocarbons and CO.

[0055] If a source of molecular oxygen, for example molecular oxygen with a purity of ≧90% by volume and ≦10% by volume of N2, is added to the product gas mixture 1 leaving the first reaction stage to produce the initial reaction gas mixture 2, preferably any of the mixing devices described above are applied accordingly (in such a case the flow of product gas mixture 1 will have the position that flow Z has above).

[0056] In the transition from a first to a second steady-state operating condition, in this case in the second steady-state operating condition, the loading of the initial reaction gas mixture 1 into the fixed catalyst bed in the first reaction stage is higher under otherwise unchanged conditions than in the first steady-state operating condition, and for reasons of safety (always outside the explosive region during the transition) and other reasons favorable to proceed in this manner, the propene load is increased many times in small successive steps, and after each such step, before carrying out the next such step, the molecular oxygen load (including any molecular oxygen-containing secondary gas introduced between the first and second reaction stages) is correspondingly increased (equilibrium), and so on until a new steady-state operating condition is reached (both steady-state operating conditions are assumed to be non-explosive). In the second steady-state operating condition, the loading of the initial reaction gas mixture 1 into the fixed catalyst bed in the first reaction stage is, under otherwise unchanged conditions, lower than in the first steady-state operating condition, and for safety and other reasons correspondingly advantageous to proceed with the transition in such a manner, the molecular oxygen loading (including any molecular oxygen containing secondary gas introduced between the first and second reaction stages) is reduced many times in small successive steps, after each such step the propene loading is correspondingly reduced (equilibrium) before carrying out the next such step, until a new steady-state operating condition is reached (both steady-state operating conditions are assumed to be non-explosive).

[0057] To start up the production process, the individual feed streams necessary for forming the initial reaction gas mixture 1 are opened in the opposite order (reverse order) as when closing these feed streams for shutting down the production process. Instead of the recycle gas stream P, lean air can be used initially. This is a mixture of 4-6% by volume O2 and 94-96% by volume N2. The subsequent feed streams are preferably started only after the volume flow rate of the previously opened feed stream has already reached at least 70% of its target value in stationary operation.

[0058] Another option for starting the method of the invention is to proceed in a manner equivalent to that taught in EP-A1 180 508.

[0059] Furthermore, a method according to the invention is preferred in which the following conditions are simultaneously met: a) the initial reaction gas mixture 1 contains 3-10% by volume of propene, 3.9-23% by volume of molecular oxygen, and contains molecular oxygen and propene in a molar ratio of O2:propene of 1.3-2.1, and (only if acrylic acid is the target product, otherwise b) is not applicable); b) The initial reaction gas mixture 2 contains 3-10% by volume acrolein, 1.8-25% by volume molecular oxygen, and a molar ratio of O2:acrolein in the range of 0.6-2.5.

[0060] In addition, the method according to the invention is particularly preferred in which the following conditions are simultaneously met: a) the initial reaction gas mixture 1 contains 4-9% by volume of propene, 5.6-19.8% by volume of molecular oxygen, and contains molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.4-2.2, and (only if acrylic acid is the target product, otherwise b) is not applicable); b) The initial reaction gas mixture 2 contains 4-9% by volume acrolein, 2.8-18% by volume molecular oxygen, and a molar ratio of O2:acrolein in the range of 0.7-2.0.

[0061] Even more preferred is a method according to the invention in which the following conditions are simultaneously met: a) the initial reaction gas mixture 1 contains 5-8% by volume of propene, 7.5-16.8% by volume of molecular oxygen, and contains molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.5-2.1, and (only if acrylic acid is the target product, otherwise b) is not applicable); b) The initial reaction gas mixture 2 contains 4.5-8% by volume acrolein, 3.6-12% by volume molecular oxygen, and a molar ratio of O2:acrolein in the range of 0.8-1.5.

[0062] As already mentioned above, the initial reaction gas mixture 1 of the process according to the invention is essentially formed (prepared) from at least three feed streams: a) fresh propene (= propene not yet used in the process; propene feedstock), e.g. refinery propene, chemical grade propene and / or polymer grade propene-containing streams; b) fresh molecular oxygen (= molecular oxygen not yet used in the process; source of molecular oxygen), for example a stream containing molecular oxygen with a purity of at least (≧) 95% by volume (relative to the total amount) O2 and not more than (≦) 5% by volume N2, and c) Recycle gas stream P.

[0063] The further feed stream for the formation of the initial reaction gas mixture 1 may be at least one hydrocarbon stream selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane and ethene (preferably consisting of methane, natural gas and propane).

[0064] Furthermore, a stream of water vapor (H2O) may be a feed stream for the formation of the reaction gas mixture 1. Such a feed stream of water vapor is not always preferred, since the separation of the target product from the water in the separation zone (especially in the case of the target product acrylic acid) is energy consuming.

[0065] In this document, the loading of a reaction gas (e.g. starting reaction gas mixture 1) into a fixed catalyst bed is taken to mean the amount of reaction gas in standard liters that is taken up through 1 liter of fixed catalyst bed per hour (=l(STP); the volume in liters that the corresponding amount of reaction gas takes up under standard conditions (0 °C, 1 bar)). The loading can also be based on only one component of the reaction gas. In that case, the loading is the amount in standard liters of this component that is taken up as a constituent of the reaction gas through 1 liter of fixed catalyst bed per hour (l(STP) / l·h). Pure inert material beds (bed sections) are not counted here against the fixed catalyst bed, whereas fixed beds (fixed bed sections) consisting of a mixture of catalyst and inert shaped diluent bodies are included (counted) in their entirety here.

[0066] In this document, the term synthesis gas (syngas) refers to a gas that is a mixture of molecular hydrogen (H2) and at least one of the carbon oxides CO (carbon monoxide) and CO2 (carbon dioxide), and optionally, hydrocarbons C n H m (e.g. methane, ethane, propane, n-butane, isobutane and ethene) (where 1≦n≦4 and 4≦m≦10), molecular nitrogen (N2), molecular oxygen (O2), water (H2O) and the noble gases (in particular He, Ne, Ar, Kr, Xe and Rn), with the proviso that the total molar amount of H2, CO and CO2 contained in gas mixture G relative to the total molar amount of all constituents contained in gas mixture G is at least 90 mol-%; the total molar amount of H2, CO, CO2 and constituents from the C* group contained in the gas mixture G relative to the total molar amount of all constituents contained in the gas mixture G is at least 99 mol %; the total molar amount of CO and CO2 contained in gas mixture G is at least 25 mol % but not more than 50 mol % relative to the total molar amount of H2, CO and CO2 contained in gas mixture G This is subject to the condition that:

[0067] Such gases are named "synthesis gas" since they can be used in a manner known per se for the synthesis of chemical compounds containing C and H and optionally O (for example, US-A 2009 / 0018220 and WO 2010 / 067945 describe the synthesis of methanol from synthesis gas containing CO and CO2; WO 2015 / 145311 describes the synthesis of hydrocarbons by Fischer Tropsch synthesis from synthesis gas containing CO and CO2; furthermore, synthesis gas can be used, for example, for the production of oxygenates via the hydroformylation of unsaturated hydrocarbons, such as olefins and acetylenes; WO No. 2005 / 095315 discloses the hydroformylation of propene to butyraldehyde using CO-rich synthesis gas; the butyraldehyde can be further converted to butanol by hydrogenation as described in EP-B 1027315 or US-A 3925490; US-A 5041683 describes the synthesis of alcohols and aldehydes from alkenes and synthesis gas; butanol and acrylic acid (optionally obtained by the process of the present invention) can be synthesized in US-A 288659 1 to butyl acrylate; US-A 5,227,544 describes the synthesis of 2-ethylhexanol from butyraldehyde (which can be esterified with acrylic acid to produce 2-ethylhexyl acrylate); acetic acid (2CO+2H→CHCOOH), ethanol (CHOH+CO+H→CH-CH-OH) and glycol (2CO+3H→HOCH-CHOH) can also be produced from synthesis gas).

[0068] It is well known that synthesis gas can be produced, for example, by (see, for example, Ullmann's Encyclopedia of Industrial Chemistry, Fifth Completely Revised Edition, Volume A12 (Formamides to Hexamethylendiamine), VCH Verlagsgesellschaft mbH, D-6940 Weinheim, pages 178 and 179, section 1.3.) (usually such production is deliberate, i.e. the synthesis gas produced is directly processed to obtain further products): -Heterogeneously catalyzed steam reforming of hydrocarbons in the gas phase: C x H y +xH2O⇔xCO+(x+0.5y)H2, and -Heterogeneously catalyzed hydrogenation of CO2 in the gas phase: CO2+H2⇔CO+H2O

[0069] Both reactions are preferably carried out at high temperature and pressure (see, for example, US-6312658, US-A4233179 and WO 2009 / 000494) and with the application of known catalysts. The combined application of both reactions (see, for example, WO 2015 / 135968 and US-B4233179) makes it possible to control the composition of the resulting product gas mixture ("synthesis gas") within certain limits as desired.

[0070] According to US-A 3,506,418, gaseous hydrocarbons can also be reformed by drying, using carbon dioxide instead of water vapor, for example: 3CO2+C3H8⇔6CO+4H2, or CO2+CH4⇔2CO+2H2

[0071] Through partial gas-phase combustion of hydrocarbons and / or organic compounds composed of C, H and O, the formation of CO and HO is possible: C xH y +(x / 2+y / 4)O2⇔xCO+y / 2H2O

[0072] In the case of a gas mixture containing all relevant reactants (as is the case for certain off-gases of the method of the invention), the overlap of all four reaction types outlined above can be applied to produce the intended synthesis gas composition. However, it is also possible to add one or more of these reactants to the off-gas produced by the method of the invention before using it for the production of any synthesis gas, for example one or more of the hydrocarbons from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane and ethene. However, preferably, such hydrocarbons (preferably propane and / or methane (respectively natural gas)) are already part of the initial reaction gas mixture 1, their presence being due to their own combustibility as well as their molar heat capacity C which allows these compounds to absorb the heat of reaction to a relatively high extent. p The suppression of the formation of undesired hot spots in the fixed catalyst bed by increasing the ignition temperature has a positive effect on the ignition behavior of the initial reaction gas mixture 1 as a whole (see, for example, US Pat. No. 6,090,977 and US Pat. No. 5,684,188).

[0073] One of the main aspects of the present invention is that the associated off-gas released from the process carried out for the production of the target products acrolein or acrylic acid has only a limited content of molecular nitrogen. This is because such molecular nitrogen is ultimately contained in the "synthesis gas" produced. However, chemical synthesis processes based on synthesis gas (e.g., hydroformylation of propene) are usually continuous processes that do not result in 100% conversion of the starting materials involved (e.g., propene). Thus, the product mixture produced still contains unreacted feedstock and inert molecular nitrogen. It is of course economically desirable for said unreacted feedstock to be recycled to the respective synthesis process. However, for a continuous chemical synthesis gas process (e.g., hydroformylation) to be successfully operated on an industrial scale, it is important to establish a steady state and degree of reaction between the feed materials, including any recycle. It is therefore important to prevent the accumulation (build-up) of excess N2 in the reaction system by recycling unreacted components. However, molecular nitrogen is difficult to separate from other compounds, such as CO, CO2, propene, H2 and others involved. Therefore, to prevent significant N2 accumulation due to recycle, it is necessary to vent at least a portion of the molecular nitrogen. The higher the N2 content of the off-gas used for synthesis gas production, the more significant this removal becomes. Since the removal of N2 necessarily involves the removal of some residual feedstock, this automatically introduces some undesirable inefficiencies and economic liabilities into the process.

[0074] In this regard, it is advantageous that due to the increase in volume (corresponding to the increase in the number of molecules) during the production of synthesis gas from the off-gas, the concentration of molecular nitrogen in the transition from off-gas to synthesis gas can drop sharply to a level that is still economically rather acceptable for the purposes of the subsequent use of the synthesis gas produced.

[0075] The realization of a heterogeneously catalyzed partial oxidation of propene to the target product acrolein (and usually small amounts of acrylic acid as a by-product) in one stage or a heterogeneously catalyzed partial oxidation of propene to the target product acrylic acid (and usually small amounts of acrolein as a by-product) in two stages using an initial reaction gas mixture 1 is described for similar reaction gas mixtures in the documents EP-A 700714 (first reaction stage; as described and also corresponding initial reaction gas mixture 1 on salt bath and tube bundle reactors in countercurrent mode), EP-A 700893 (second reaction stage; as described and also corresponding initial reaction gas mixture 2 on salt bath and tube bundle reactors in countercurrent mode), WO 04 / 085369 (two-stage process), WO 04 / 85363, DE-A10313212 (first reaction step), EP-A1159248 (two-step process), EP-A1159246 (second reaction step), EP-A1159247 (two-step process), DE-A19948248 (two-step process), WO04 / 085368 (two-step process), DE-A102004021764 (two-step process), WO 04 / 085362 (first reaction step), WO 04 / 085370 (second reaction step), WO 04 / 085365 (second reaction step), WO This can be carried out in a manner equivalent to that described in EP-A 04 / 085367 (two-stage process), EP-A 990636, EP-A 1007007 and EP-A 1106598.

[0076] This applies in particular to all the examples contained in these documents. These can be carried out as described in these documents, with the difference that the starting reaction gas mixture used for the first reaction stage (propene to acrolein) is the starting reaction gas mixture 1 produced according to the invention. With respect to the remaining parameters, the procedure is as in the examples of the described documents (in particular with regard to the fixed catalyst bed and the loading of the reactants of the fixed catalyst bed). If the procedure of the above-mentioned examples of the prior art is of two stages and there is secondary oxygen fed between the two reaction stages, the feeding is carried out in a corresponding manner, but with the amount adjusted so that the molar ratio of molecular oxygen to acrolein in the starting reaction gas mixture 2 of the second reaction stage corresponds, if necessary, to the molar ratio in the documents (in the range from 0.5 to 3).

[0077] Advantageously according to the invention, the amount of molecular oxygen in the initial reaction gas mixture 1, respectively in the initial reaction gas mixture 2 is such that the product gas mixture 1, respectively in the second reaction stage on leaving the first reaction stage still comprises unconverted molecular oxygen (suitably ≧0.5-6 vol.%, advantageously 1-5 vol.%, preferably 2-4 vol.%).

[0078] Multimetal oxide catalysts particularly suitable for the respective reaction steps have been described many times before and are well known to those skilled in the art, for example EP-A 253 409 refers to page 5 of the corresponding US patent.

[0079] Preferred catalysts for the respective oxidation step are also disclosed by DE-A 4 431 957, DE-A 10 2004 025 445 and DE-A 4 431 949. This applies in particular to the general formula I of the two of these three mentioned documents which have the earlier publication date. Particularly advantageous catalysts for the respective oxidation step are disclosed in documents DE-A 10 325 488, DE-A 10 325 487, DE-A 10 353 954, DE-A 10 344 149, DE-A 10 351 269, DE-A 10 350 812 and DE-A 10 350 822.

[0080] Applicable catalysts for the first reaction step of the process of the present invention are those whose active material is at least one multimetal oxide composition containing Mo, Bi and Fe.

[0081] Such catalytically active multimetal oxide compositions are in particular the multimetal oxide active compositions of the general formula I of DE-A 19955176, the multimetal oxide active compositions of the general formula I of DE-A 19948523, the multimetal oxide active compositions of the general formulae I, II and III of DE-A 10101695, the multimetal oxide active compositions of the general formulae I, II and III of DE-A 19948248 and the multimetal oxide active compositions of the general formulae I, II and III of DE-A 19955168, and furthermore the multimetal oxide active compositions specified in EP-A 700714.

[0082] Suitable for the first reaction step are also multimetal oxide catalysts containing Mo, Bi and Fe, which are disclosed in the documents Research Disclosure No. 497012, dated 29.08.2005, DE-A 10046957, DE-A 10063162, DE-C 3338380, DE-A 19902562, EP-A 15565, DE-C 2380765, EP-A 807465, EP-A 279374, DE-A 3300044, EP-A 575897, US-A 4438217, DE-A 19855913, WO 98 / 24746, DE-A 19746210 (of general formula II), JP-A 91 / 294239, EP-A 293224 and EP-A 700714. This applies in particular to the exemplary embodiments in these documents, among which EP-A 15565, EP-A 575897, DE-A 19746210 and DE-A 19855913 are preferred. In this context, catalysts according to Example 1c of EP-A 15565 and active compositions prepared in a corresponding manner, the active compositions of which contain Mo 12 Ni 6.5 Zn2Fe2Bi1P 0.0065 K 0.06 O xThe catalyst with the composition Mo. 10SiO2 is of interest. 12 Co7Fe3Bi 0.6 K 0.08 S 1.6 O x Attention is also drawn to catalysts as unsupported hollow cylinder catalysts of the shape 5 mm x 3 mm x 2 mm (external diameter x height x internal diameter) and also to the unsupported multimetal oxide II catalyst according to example 1 of DE-A 19746210. Mention should also be made of the multimetal oxide catalysts of US-A 4438217. This multimetal oxide catalyst is particularly suitable when these hollow cylinders have the shape 5.5 mm x 3 mm x 3.5 mm, or 5 mm x 2 mm x 2 mm, or 5 mm x 3 mm x 2 mm, or 6 mm x 3 mm x 3 mm, or 7 mm x 3 mm x 4 mm (respectively external diameter x height x internal diameter). Further possible catalyst shapes in this context are extrudates (for example length 7.7 mm and diameter 7 mm; or length 6.4 mm and diameter 5.7 mm).

[0083] A number of multimetal oxide active compositions suitable for the propene to acrolein process can be encompassed by the general formula I. Mo 12 Bi a Fe b X 1 c X 2 d X 3 e X 4 f O n (I) wherein the variables are defined as follows: X 1 = nickel and / or cobalt, X 2 = thallium, alkali metals and / or alkaline earth metals, X 3 = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten, X 4 = silicon, aluminium, titanium and / or zirconium, a=0.5~5, b=0.01 to 5, preferably 2 to 4, c=0 to 10, preferably 3 to 10, d=0 to 2, preferably 0.02 to 2; e=0 to 8, preferably 0 to 5; f=0~10 and n = a number determined by the valence and frequency of the elements other than oxygen in I.

[0084] These are available in a manner known per se (see, for example, DE-A 4023239) and are usually formed from substances into spheres, rings or cylinders or are otherwise used in the form of coated catalysts, i.e. preformed inert support bodies which are coated with the active composition.

[0085] In principle, the active compositions of general formula I can be prepared in a simple manner by obtaining a finely divided dry mixture with a composition corresponding to their stoichiometry, preferably from suitable sources of these elemental components, and calcining it at a temperature between 350 and 650° C. Calcination can be carried out either under inert gas or under an oxidizing atmosphere, for example air (a mixture of inert gas and oxygen), or even under a reducing atmosphere (for example a mixture of inert gas, NH3, CO and / or H2). The calcination time can be from a few minutes to a few hours and usually decreases with temperature. Useful sources for the elemental components of the multimetal oxide active compositions I are compounds which are already oxides and / or which can be converted to oxides by heating, at least in the presence of oxygen.

[0086] In addition to the oxides, such useful starting compounds include, inter alia, halides, nitrates, formates, oxalates, citrates, acetates, carbonates, amine complexes, ammonium salts and / or hydroxides (compounds which decompose and / or can decompose at the latest after calcination, such as, for example, NH4OH, (NH4)2CO3, NH4NO3, NH4CHO2, CH3COOH, NH4CH3CO2 and / or ammonium oxalate, can further be incorporated into the intimate dry mixture in order to obtain compounds which can be released in gaseous form).

[0087] The starting compounds for preparing the multimetal oxide active composition I can be intimately mixed in dry or wet form. If they are mixed in dry form, the starting compounds are suitably used as finely divided powders and, after mixing and optional pressing, are subjected to calcination. However, intimate mixing in wet form is preferred. Usually, the starting compounds are mixed with one another in the form of an aqueous solution and / or suspension. Intimate dry mixtures are obtained in the described mixing process, especially when the starting materials represent exclusively the source of elemental components in dissolved form. The solvent used is preferably water. The aqueous composition obtained is subsequently dried, the drying process being preferably carried out by spray drying the aqueous mixture at an outlet temperature of 100-150°C.

[0088] The active compositions of multimetal oxides of general formula I are used for the "propene → acrolein" process formed into a certain catalyst shape, which can be carried out either before or after the final calcination. For example, unsupported catalysts can be prepared from the active compositions in powder form or their uncalcined and / or partially calcined precursor compositions by compressing them into the desired catalyst shape (for example by tableting or extrusion), optionally with the addition of auxiliaries such as graphite or stearic acid as lubricants and / or forming auxiliaries and reinforcing agents such as glass microfibers, asbestos, silicon carbide or potassium titanate. Instead of graphite, it is also possible to use hexagonal boron nitride as an auxiliary for shaping, as recommended in DE-A102005037678. Examples of suitable unsupported catalyst shapes include solid or hollow cylinders with an outer diameter and length of 2 to 10 mm. In the case of hollow cylinders, a wall thickness of 1 to 3 mm is advantageous. The unsupported catalyst may also naturally have a spherical shape, the diameter of the sphere being between 2 and 10 mm.

[0089] A particularly preferred hollow cylinder shape is 5 mm x 3 mm x 2 mm (external diameter x length x internal diameter), especially for unsupported catalysts.

[0090] Naturally, the uncalcined and / or partially calcined powdered active composition or its powdered precursor composition can also be shaped by coating a previously shaped inert catalyst support body. The coating of the support body to produce the coated catalyst is generally carried out in a suitable rotatable vessel, as disclosed, for example, in DE-A 2909671, EP-A 293859 or EP-A 714700. To coat the support body, the powder composition to be applied is appropriately moistened and, after coating, is dried again, for example by means of hot air. The coating thickness of the powder composition applied to the support body is appropriately selected in the range of 10 to 1000 μm, preferably in the range of 50 to 500 μm, more preferably in the range of 150 to 250 μm.

[0091] Useful support materials are conventional porous or non-porous aluminum oxide, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates, such as magnesium silicate or aluminum silicate. These generally behave substantially inert with respect to the target reactants on which the method according to the invention is based. The support body can have a regular or irregular shape, but regularly shaped support bodies with a well-defined surface roughness are preferred, for example spheres or hollow cylinders. Suitable support bodies are substantially non-porous, surface-roughened spherical supports made of steatite with a diameter of 1-10 mm or 1-8 mm, preferably 4-5 mm. However, suitable support bodies are also cylinders with a length of 2-10 mm and an external diameter of 4-10 mm. In the case of rings suitable as support bodies according to the invention, the wall thickness is also usually 1-4 mm. Annular support bodies to be used preferentially according to the invention have a length of 2-6 mm, an external diameter of 4-8 mm and a wall thickness of 1-2 mm. A suitable support body according to the invention is in particular also a ring of the shape 7 mm x 3 mm x 4 mm (external diameter x length x internal diameter). The fineness of the catalytically active oxide composition applied to the surface of the support body is naturally adjusted to the desired coating thickness (see EP-A 714 700).

[0092] The multimetal oxide active compositions to be preferably used in the propene to acrolein process are also of the general formula II [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" (II) wherein the variables are defined as follows: Z2 = molybdenum or molybdenum and tungsten, Z 3 = nickel and / or cobalt, Z 4 = thallium, alkali metals and / or alkaline earth metals, Z 5 = phosphorus, arsenic, boron, antimony, tin, cerium and / or lead, Z 6 = silicon, aluminium, titanium and / or zirconium, Z 7 = copper, silver and / or gold, a"=0.1~1, b"=0.2~2, c"=3~10, d"=0.02~2, e"=0.01 to 5, preferably 0.1 to 3, f"=0~5, g"=0~10, h"=0~1, x", y" = a number determined by the valence and frequency of elements other than oxygen in II, p" and q" are numbers whose p" / q" ratio is 0.1 to 5, preferably 0.5 to 2. and are separated from these local environments because their composition differs from them, and their maximum diameter (the longest direct line passing through the center of the region and joining two points on the surface (boundary) of the region) is between 1 nm and 100 μm, often between 10 nm and 500 nm or between 1 μm and 50 or 25 μm. a" Z 2 b" O x" The composition also includes a three-dimensional domain of

[0093] Z 2 b" = (tungsten) b" and Z 2 12 = (Molybdenum) 12 Very particular preference is given to composition II in which:

[0094] The total proportion of multimetal oxide composition II [Bi a" Z 2 b" O x" ] p" At least 25 mol % (preferably at least 50 mol %, more preferably at least 100 mol %) of the compound having the chemical composition [Bi a" Z 2 b" O x" It is also advantageous if the three-dimensional domains are in the form of three-dimensional domains of the formula

[0023] , which domains are separated from their local environments because their chemical composition differs from these local environments and whose maximum diameter is in the range of 1 nm to 100 μm.

[0095] With regard to shaping, the statements made for the multimetal oxide composition I catalysts apply correspondingly to the multimetal oxide composition II catalysts.

[0096] The preparation of multimetal oxide active compositions II is described, for example, in EP-A 575897 and DE-A 19855913.

[0097] The inert support materials recommended above are also useful, inter alia, as inert materials for diluting and / or partitioning suitable fixed catalyst beds or as pre-beds for protecting them and / or heating the gas mixture.

[0098] For the second step (second reaction stage), the heterogeneously catalyzed gas-phase partial oxidation of acrolein to acrylic acid, useful active multimetal oxide compositions for the required catalysts have already been mentioned, in principle all multimetal oxide compositions containing Mo and V, as is the case, for example, in DE-A 10046928.

[0099] Many of them, for example those of DE-A 19815281, can be encompassed by the general formula III Mo 12 V a X 1 b X 2 c X 3d X 4 e X 5 f X 6 g O n (III) wherein the variables are 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=1~6, b=0.2~4, c=0.5~18, d=0~40, e=0~2, f=0~4, g=0~40 and n = a number determined by the valence and frequency of the elements other than oxygen in III.

[0100] Among the active multimetal oxides III, preferred embodiments according to the invention are encompassed by the following definitions of the variables of general formula III: 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=1.5~5, b=0.5~2, c=0.5~3, d=0~2, e=0~0.2, f=0~1 and n = a number determined by the valence and frequency of the elements other than oxygen in III.

[0101] However, very particularly preferred multimetal oxides III according to the invention are those of the general formula IV Mo 12 V a' Y 1 b' Y 2 c' Y 5 f' Y 6 g' O n' (IV) [In the formula, 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'=2~4, b'=1~1.5, c'=1~3, f'=0~0.5, g'=0~8 and n' = a number determined by the valence and frequency of the elements other than oxygen in IV.

[0102] Multimetal oxide active compositions (III) suitable according to the invention are available in a manner known per se, for example as disclosed in DE-A 4 335 973 or EP-A 714 700.

[0103] In principle, multimetal oxide active compositions suitable for the "acrolein → acrylic acid" process, in particular compositions of general formula III, can be prepared in a simple manner by obtaining a very intimate, preferably finely divided, dry mixture with a composition corresponding to their stoichiometry from suitable sources of these elemental components and calcining it at temperatures between 350 and 600 ° C. Calcination can be carried out either under inert gas or under an oxidizing atmosphere, for example air (a mixture of inert gas and oxygen), and alternatively under a reducing atmosphere (for example a mixture of inert gas and reducing gas, for example H2, NH3, CO, methane and / or acrolein or the reducing gases themselves as described). Calcination times can be from a few minutes to a few hours and usually decrease with temperature. Useful sources for the elemental components of the multimetal oxide active composition III include compounds which are already oxides and / or which can be converted to oxides by heating, at least in the presence of oxygen.

[0104] The starting compounds for the preparation of the multimetal oxide composition III can be intimately mixed in dry or wet form. If they are mixed in dry form, the starting compounds are suitably used in the form of finely divided powders, which are mixed and, if appropriate, compressed and then calcined. Intimate mixing in wet form is, however, preferred.

[0105] This is usually done by mixing the starting compounds with one another in the form of an aqueous solution and / or suspension. In cases where the starting materials represent exclusively the source of elemental components in dissolved form, a particularly intimate dry mixture is obtained in the described mixing process. The solvent used is preferably water. The aqueous composition obtained is then dried, the drying process being preferably carried out by spray drying the aqueous mixture at an outlet temperature of 100-150°C.

[0106] The resulting multimetal oxide compositions, in particular those of general formula III, used for acrolein oxidation are formed into a certain catalyst shape, which can be carried out before or after the final calcination. For example, unsupported catalysts can be prepared from the active composition in powder form or its uncalcined precursor composition by compressing it into the desired catalyst shape (for example by tableting or extrusion), optionally with the addition of auxiliaries such as graphite or stearic acid as lubricants and / or forming auxiliaries and reinforcing agents such as glass microfibers, asbestos, silicon carbide or potassium titanate. Examples of suitable unsupported catalyst shapes are solid or hollow cylinders with an outer diameter and length of 2 to 10 mm. In the case of hollow cylinders, a wall thickness of 1 to 3 mm is appropriate. Unsupported catalysts can of course also have a spherical shape, in which case the diameter of the spheres can be 2 to 10 mm (for example 8.2 mm or 5.1 mm).

[0107] The uncalcined, powdered active composition or its powdered precursor composition can of course also be shaped by application to a preformed inert catalyst support. The coating of the support body to prepare the coated catalyst is generally carried out in a suitable rotatable vessel, as disclosed, for example, in DE-A 2909671, EP-A 293859 or EP-A 714700.

[0108] To coat the support body, the powder composition to be applied is appropriately moistened and, after application, is dried again, for example, by means of hot air. The coating thickness of the powder composition to be applied to the support body is appropriately selected within the range of 10 to 1000 μm, preferably within the range of 50 to 500 μm, more preferably within the range of 150 to 250 μm.

[0109] Useful support materials are conventional porous or non-porous aluminum oxide, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates, such as magnesium silicate or aluminum silicate. The support body can have a regular or irregular shape, but a regularly shaped support body with a defined surface roughness is preferred, for example a sphere or a hollow cylinder with a dust layer. Suitable support bodies include substantially non-porous, surface-roughened, spherical supports made of steatite, the diameter of which is 1-10 mm or 1-8 mm, preferably 4-5 mm. In other words, suitable spherical shapes may have a diameter of 8.2 mm or 5.1 mm. However, suitable support bodies also include cylinders with a length of 2-10 mm and an outer diameter of 4-10 mm. In the case of a ring as a support body, the wall thickness is also usually 1-4 mm. The annular supports to be used preferentially have a length of 2 to 6 mm, an outer diameter of 4 to 8 mm and a wall thickness of 1 to 2 mm. Suitable support bodies are also in particular rings of the shape 7 mm x 3 mm x 4 mm (external diameter x length x internal diameter). The fineness of the catalytically active oxide composition applied to the surface of the support body is of course adapted to the desired coating thickness (see EP-A 714 700).

[0110] Particularly suitable multimetal oxide catalysts for the "acrolein→acrylic acid" process are also the multimetal oxide catalysts of DE-A 19815281, in particular if they have the active composition of multimetal oxides of the general formula I of this document.

[0111] Advantageously, unsupported catalyst rings are used in the coated catalyst rings for the propene to acrolein and acrolein to acrylic acid processes.

[0112] The partial oxidation of propene to acrolein according to the invention can be carried out with the catalysts described, for example, in a single-zone, multiple catalyst tube fixed-bed reactor, as described in DE-A 4431957. In this case, the reaction gas mixture and the heat transfer medium (heat exchange medium) can be conducted in cocurrent or countercurrent flow, as viewed from above the reactor.

[0113] The reaction operating pressure is in the range of 1-6 bar (= absolute pressure; the highest value is at the inlet to the first reaction stage (to the corresponding reactor)), the total space velocity of the initial reaction gas mixture 1 in the first fixed catalyst bed is preferably 1500-4000 or 1500-6000 l(STP) / l·h or more. The propene loading (hourly space velocity of propene over the first fixed catalyst bed) is usually 50-200 l(STP) / l·h or 90-200 l(STP) / l·h or more than 50-300 l(STP) / l·h or more than 90-300 l(STP) / l·h. Propene loadings above 135 l(STP) / l·h, or ≥ 140 l(STP) / l·h, or ≥ 150 l(STP) / l·h, or ≥ 160 l(STP) / l·h, or ≥ 170 l(STP) / l·h are particularly preferred according to the invention, since the inventive starting reaction gas mixture 1 fed to reaction stage 1 leads, by its composition, to a favourable hot spot behaviour (all of the above applies irrespective of the particular choice of the fixed bed reactor).

[0114] The flow of the charge gas mixture (initial reaction gas mixture 1) into the single-zone multiple catalyst tube fixed bed reactor is preferably from above. The heat exchange medium used is suitably a salt melt, preferably consisting of 60% by weight potassium nitrate (KNO3) and 40% by weight sodium nitrite (NaNO2), or 53% by weight potassium nitrate (KNO3), 40% by weight sodium nitrite (NaNO2) and 7% by weight sodium nitrate (NaNO3).

[0115] The salt melt and the reaction gas mixture can, as already mentioned, be conducted either in co-current or counter-current flow as viewed from above the reactor. The salt melt itself is preferably conducted in serpentine fashion around the catalyst tubes.

[0116] If the flow of starting reaction gas mixture 1 into the catalyst tubes is from top to bottom, it is appropriate to fill the catalyst into the catalyst tubes from bottom to top as follows (for a bottom to top flow of starting reaction gas mixture 1 the filling order is appropriately reversed): - firstly, filling the catalytic tube with either the catalyst alone or a mixture of catalyst and inert material to a length of 40-80 or 40-60% of the length of the catalytic tube, the latter constituting a mass ratio of up to 30 or 20% by mass relative to the mixture (section C); - followed by filling the tube with either the catalyst alone or a mixture of catalyst and inert material to a length of 20-50 or 20-40% of the total length of the tube, the latter constituting a mass ratio of up to 40% by mass relative to the mixture (section B); and - Finally, a bed of inert material, preferably chosen to cause a minimum pressure drop, is filled over a length of 10-20% of the total tube length (section A).

[0117] Section C is preferably not diluted.

[0118] A particularly suitable variant of the abovementioned packing, when a catalyst is used, is according to Research Disclosure No. 497012, dated 29 August 2005, Example 1 of DE-A 10046957 or Example 3 of DE-A 10046957, and the inert material used is a steatite ring having the dimensions 7 mm x 7 mm x 4 mm (external diameter x height x internal diameter). With regard to the salt bath temperature, the statements in DE-A 4 431 957 apply accordingly.

[0119] However, the partial oxidation of propene to acrolein in the first reaction step can also be carried out using the described catalysts in a two-zone multiple catalyst tube fixed bed reactor, as described, for example, by DE-A19910506, DE-A102005009885, DE-A102004032129, DE-A102005013039, DE-A102005009891 and DE-A102005010111. Advantageously according to the invention, the inventive partial oxidation of propene to acrolein is carried out as described in EP-A 1 159 244, and most preferably as described in WO 2000 / 53558, WO 04 / 085363, WO 2004 / 085367, WO 2004 / 085369, WO 2004 / 085368, WO 04 / 085362, WO 2004 / 085369, WO 2007 / 08287 and WO 2012 / 049246.

[0120] In other words, the partial oxidation of propene to acrolein to be carried out in the first reaction stage can be particularly advantageously carried out on a first fixed catalyst bed having a higher propene loading and at least two temperature zones.

[0121] In this connection, reference is made to the documents EP-A1 159244, WO 2000 / 53558, WO 04 / 085363, WO 2004 / 085367, WO 2004 / 085369, WO 2004 / 085368, WO 04 / 085362, WO 2007 / 08287 and WO 2012 / 049246, which are also considered to be an integral part of this document.

[0122] The implementation of the second step in the case of the two-stage partial oxidation of propene to acrylic acid, i.e. the partial oxidation of acrolein to acrylic acid, can be carried out with the catalysts described, for example in a one-zone, multiple catalyst tube fixed bed reactor, as described in DE-A 4431949. In this reaction stage, the initial reaction gas mixture 2 and the heat transfer medium can be conducted in cocurrent or countercurrent flow, as viewed from above the reactor. In general, the product gas mixture of the preceding inventive partial oxidation of propene to acrolein is in principle conducted thus (optionally after intermediate cooling, which is carried out indirectly or directly, for example by adding secondary molecular oxygen), i.e. without prior removal of secondary components, to the second reaction stage, i.e. the partial oxidation of acrolein.

[0123] The molecular oxygen required for the second step, the partial oxidation of acrolein, may already be present in the starting reaction gas mixture 1 for the partial oxidation of propene to acrolein, but it can also be added partly or completely directly to the product gas mixture 1 of the first reaction stage, i.e. the partial oxidation of propene to acrolein, which can be carried out in the form of oxygen with a purity of ≧90% by volume O2 and ≦90% by volume N2.

[0124] As in the case of the first reaction stage (propene → acrolein), the reaction pressure in the second reaction stage (acrolein → acrylic acid) is also usually in the range of 1-6 bar (= absolute pressure; the highest value is at the inlet to the second reaction stage (to the corresponding reactor)), the total space velocity of the starting reaction gas mixture 2 in the second fixed catalyst bed is preferably 1500-4000 or 1500-6000 l(STP) / l·h or more. The acrolein loading (hourly space velocity of acrolein over the second fixed catalyst bed) is usually 40-190 l(STP) / l·h or 80-190 l(STP) / l·h or more than 40-290 l(STP) / l·h or more than 80-290 l(STP) / l·h. Acrolein loadings above 125 l(STP) / l·h, or ≧130 l(STP) / l·h, or ≧140 l(STP) / l·h, or ≧150 l(STP) / l·h, or ≧160 l(STP) / l·h are particularly preferred, since the starting reaction gas mixture 2 which is fed to the reaction stage 2 according to the invention, due to its composition, leads to a favourable hot spot behaviour (all of the above applies irrespective of the particular choice of the fixed bed reactor).

[0125] The flow of the charge gas mixture (initial reaction gas mixture 2) in the single-zone multiple catalyst tube fixed bed reactor is likewise preferably from above. The heat exchange medium used in the second stage is also suitably a salt melt, preferably consisting of 60% by weight of potassium nitrate (KNO3) and 40% by weight of sodium nitrite (NaNO2), or 53% by weight of potassium nitrate (KNO3), 40% by weight of sodium nitrite (NaNO2) and 7% by weight of sodium nitrate (NaNO3). Viewed from above the reactor, as already mentioned, the salt melt and the reaction gas mixture (charge gas mixture) can be conducted either in parallel or countercurrent. The salt melt itself is preferably conducted in a serpentine manner around the contact tubes.

[0126] If the flow of the starting reaction gas mixture 2 into the catalyst tubes is top to bottom, it is appropriate to fill the catalyst into the catalyst tubes from bottom to top as follows: - firstly, filling the catalytic tube to a length of 50-80 or 50-70% of the length with either the catalyst alone or a mixture of catalyst and inert material, the latter constituting a mass ratio of up to 30% by weight (or up to 20% by weight) relative to the mixture (section C); - followed by filling the tube with either the catalyst alone or a mixture of catalyst and inert material to a length of 20-40% of the total length of the tube, the latter constituting a mass ratio of up to 50% by mass or up to 40% by mass of the mixture (section B); and - Finally, a bed of inert material, preferably chosen to cause a minimum pressure drop, is filled over a length of 5-20% of the total tube length (section A).

[0127] Section C is preferably undiluted. As is quite common for heterogeneously catalyzed gas-phase partial oxidation of acrolein to acrylic acid (especially at high acrolein loadings on fixed catalyst beds and high water vapor content of the charge gas mixture), section B may also consist of two successive catalyst dilutions (with the aim of minimizing hot spot temperatures and their sensitivity), from bottom to top, first with up to 30 (or 20)% by weight of inert material, followed by >20% to 50 or 20% to 40% by weight of inert material. Section C is then preferably undiluted.

[0128] For a bottom-to-top flow of the initial reaction gas mixture 2 into the catalyst tubes, the filling of the catalyst tubes is appropriately reversed.

[0129] The abovementioned variations in the loadings are particularly suitable if the catalyst used is that prepared according to Example 5 of DE-A 10046928 or according to DE-A 19815281 and the inert material used is a steatite ring having the dimensions 7 mm x 7 mm x 4 mm or 7 mm x 7 mm x 3 mm (in each case external diameter x height x internal diameter). With regard to the salt bath temperature, the statements in DE-A 4431949 apply correspondingly.

[0130] However, the partial oxidation of acrolein to acrylic acid can also be carried out using the described catalysts in a two-zone multi-catalyst tube fixed-bed reactor, as described, for example, in DE-A 19910508, EP-A 1159246, WO 2004 / 085365, WO 2004 / 085370, WO 2007 / 08287, WO 2004 / 085367, WO 2004 / 085368 and WO 2004 / 085369. Also, when the acrolein partial oxidation as described above is carried out in a two-zone, multiple catalyst tube fixed bed reactor as the second stage of a two-stage propene partial oxidation to acrylic acid, the charge gas mixture (initial reaction gas mixture 2) can be suitably obtained directly by using the product gas mixture of the partial oxidation directed to the first step (propene → acrolein) (optionally after intermediate cooling, indirectly or directly, for example by feeding secondary oxygen with a purity of ≧90% by volume O2 and ≦10% by volume N2) (as already described above). The oxygen required for the acrolein partial oxidation is preferably added during the first and second reaction stages in the form of (secondary) oxygen with a purity of ≧90% by volume O2 and ≦10% by volume N2, preferably ≧95% by volume O2 and ≦5% by volume N2, or ≧98% by volume O2 and ≦2% by volume N2, or ≧99% by volume O2 and ≦1% by volume N2, or 99.4% to 99.9% by volume O2 and 0.1 to 0.6% by volume N2, or 99.5% to 99.8% by volume O2 and 0.2 to 0.5% by volume N2, for example directly to the product gas mixture 1 of the first step (propene → acrolein) of the two-stage partial oxidation. However, oxygen is also already present in the starting reaction gas mixture 1 for the first reaction stage, as described above.

[0131] In the two-stage partial oxidation of propene to acrylic acid with further direct use of the product gas mixture 1 of the first stage of partial oxidation to charge the second stage of partial oxidation, two one-zone multiple catalyst tube fixed bed reactors (at high reactant loading on the catalyst bed, as is quite common, countercurrent mode between the reactant gas and the salt bath (heat carrier) is preferred, viewed from above the tube bundle reactor) or two two-zone multiple catalyst tube fixed bed reactors are generally connected in series. A mixed series connection (one zone / two zones or vice versa) is also possible.

[0132] Advantageously, according to the invention, the inventive partial oxidation of acrolein to acrylic acid is carried out as described in DE-A 19910508, EP-A 1 159246, WO 2000 / 53558, WO 2004 / 085365, WO 2004 / 085370, WO 2007 / 08287, WO 2004 / 085367, WO 2004 / 085368 and WO 2004 / 085369.

[0133] In other words, the partial oxidation of acrolein to acrylic acid to be carried out in the second reaction stage can be particularly advantageously carried out on a second fixed catalyst bed having a higher acrolein loading and at least two temperature zones.

[0134] In this connection reference is made to documents DE-A 19910508, EP-A 1 159 246, WO 2000 / 53558, WO 2004 / 085365, WO 2004 / 085370, WO 2007 / 08287, WO 2004 / 085367, WO 2004 / 085368 and WO 2004 / 085369, which are also considered to be an integral part of this document.

[0135] Between the reactors, an intermediate cooler may be placed, which may optionally contain an inert bed capable of performing a filtering function. The salt bath temperature of the multiple catalyst tube reactor for the first step of the two-stage partial oxidation of propene to acrylic acid is generally between 250 and 500 ° C. The salt bath temperature of the multiple catalyst tube reactor for the second step of the partial oxidation of propene to acrylic acid is usually between 200 and 450 ° C. The reaction temperature is regulated by means of the salt bath temperature. In addition, the heat exchange medium (preferably salt melt) is usually led through the associated multiple catalyst tube fixed bed reactors in such an amount that the difference between their input temperature and their output temperature is generally ≦ 5 ° C (countercurrent as well as cocurrent operation). As already described, both steps of the partial oxidation of propene to acrylic acid can also be introduced in one two-zone reactor ("single reactor") on one fixed catalyst bed whose composition of multimetal oxides changes along the reaction tube from one containing Mo, Bi and Fe to one containing Mo and V. The salt bath temperature of each reaction zone is individually regulated.

[0136] Overall, a tube bundle reactor in which the catalyst loading is appropriately varied along the individual catalyst tubes forms the simplest embodiment of the two oxidation stages for the two steps of the partial oxidation of propene to acrylic acid, with completion of the first reaction step (such two-stage propene partial oxidation in a single reactor is taught, for example, by EP-A 911313, EP-A 979813, EP-A 990636 and DE-A 2830765). Optionally, the loading of the catalyst in the catalyst tubes is interrupted by an inert material bed.

[0137] However, the two oxidation stages are preferably carried out in the form of two reactors connected in series, for example two thermoplate reactors connected in series or two tube bundle reactors connected in series. The latter may also be arranged in one reactor, in which case the transition from one tube bundle type to the other occurs by a bed of inert material not contained in the catalyst tubes (and suitably readily available). The catalyst tubes are generally flowed by a heat transfer medium, which does not reach the contained inert material bed as described above. Therefore, the two catalyst tube bundle types are advantageously housed in spatially separate reactors. Generally, an intermediate cooler is placed between the two tube bundle reactors to reduce any acrolein post-combustion events in the product gas mixture 1 leaving the first oxidation stage (reactor).

[0138] Typically, for each of the two reaction stages in the reaction tubes of each tube bundle reactor, the fixed catalyst beds are packed in such a manner and the reactor design is so designed that, during production operation, if the temperature of each salt bath increases at any time and at every location in each catalyst bed by +1°C, the temperature increase in the catalyst bed is ≦+9°C, preferably ≦+7°C, better still ≦+5°C or ≦+3°C, but ≧0°C (see also, for example, US Pat. No. 8,618,336).

[0139] In the method according to the invention, for example, the metering of a cold oxygen source into the product gas mixture 1 of the first partial oxidation stage can also cause its cooling by a direct route before it is further used as a constituent of the starting reaction gas mixture 2 for the second partial oxidation stage.

[0140] At this point, it should be noted that the Applicant has surprisingly found that the catalysts applied in the two reaction stages of the method of the invention are obviously not capable of catalyzing the combustion of the by-product CO to produce CO using O2 contained in the reaction gas mixture. This fact makes it possible in the method of the invention to enrich CO by means of a relatively high recycle gas ratio K (K=(mass flow rate of the recycle gas stream P) / (total mass flow rate M of the residual gas mixture R continuously led off from the separation zone)) and to produce a CO-rich off-gas stream O. Furthermore, the use of such a CO-rich off-gas for the production of synthesis gas can be carried out in a particularly simple manner by simply mixing the off-gas as it is with synthesis gas produced elsewhere and utilizing the mixture produced as fresh synthesis gas for synthesis purposes.

[0141] In this document, the conversion U of compound X in the reaction zone X (Mole %) refers to the percentage of the molar amount of said compound X that is fed to the relevant reaction zone and that is converted to target products and by-products within the reaction zone.

[0142] In this document, the selectivity S of the formation of compound Y in the reaction zone by the conversion of compound X in said reaction zone is Y (mol %) means the molar amount of compound Y formed in the reaction zone expressed as a percentage of the molar amount of compound X converted in the reaction zone.

[0143] In this document, the yield A of product Y achieved in said reaction zone by conversion of compound X in said reaction zone is Y (mol%) is the product [U X (mol%) xS Y (mol %) divided by 100 mol %.

[0144] The product gas mixture 1, respectively the product gas mixture 2, contains the target product acrolein (with small amounts of the by-product acrylic acid), respectively acrylic acid (with small amounts of the by-product acrolein), water vapor (formed as a by-product and / or further used as inert diluent gas), any unconverted molecular oxygen (as already mentioned herein, with a view to the life of the catalysts used, it is preferred if the content of molecular oxygen contained in the respective product gas mixture is still >0% by volume, for example 1.5% to 4% by volume), as well as further pure compounds which, at an operating pressure of 1 bar, have a higher boiling point than the target products ("high boiling solvents"), such as CO, CO2, etc. ), or other by-products or secondary components having a lower boiling point ("low boiling solvents"), lower aldehydes and / or lower alkene carboxylic acids (e.g. formaldehyde, acetaldehyde, propionaldehyde, formic acid, acetic acid and propionic acid), maleic anhydride, benzaldehyde, aromatic carboxylic acids and aromatic carboxylic acid anhydrides, e.g. phthalic anhydride and benzoic acid, propane, which is further used as a leading companion to propene or as an inert diluent gas, and optionally other inert diluent gases, e.g. N2 and one or more hydrocarbons from the group consisting of methane, ethane, n-butane, isobutane and ethene, which are used simultaneously.

[0145] The separation of the respective target product (acrolein or acrylic acid) from the respective product gas mixture resulting in the respective heterogeneously catalyzed partial gas phase oxidation (from product gas mixture 1 or product gas mixture 2) is carried out in the process of the invention in a corresponding separation zone by transferring at least 95 mol %, preferably at least 96 mol %, more preferably at least 97 mol %, better still at least 98 mol % and in the best case at least 99 mol %, but usually not more than 99.8 mol %, of the respective target product contained in the respective product gas mixture from the gas phase into the liquid phase. This is carried out as known from the processes of the prior art. Generally In these methods, the product gas mixture before each entering the separation zone is first subjected to indirect and / or direct cooling (such cooling may already result in the separation of constituents of the product gas mixture continuously transferred to the separation zone containing the target product, which constituents, as pure compounds, have a boiling point at an operating pressure of 1 bar higher than the boiling point of the target product from said product gas mixture). As in the prior art, the transfer of the target product present in the product gas mixture that has been cooled in this way or present in the uncooled product gas mixture to the liquid phase is preferably by absorption in a liquid absorbent and / or by condensation (preferably fractional condensation), for example with WO The process is carried out as described in the respective prior art references cited in WO 2007 / 074045 and WO 2007 / 060036. Useful absorbents for substantially both target products are either water or aqueous solutions or organic solvents with high boiling points. Subsequent work-up of the resulting liquid phase containing the desired target product in the separation zone through application of thermal separation processes, such as stripping, distillation, rectification (including azeotropic rectification, respectively distillation), extraction and / or crystallization (all these separation processes are part of and belong to the separation zone) finally produces the target product in its desired purity, which usually leaves the separation (separating) zone continuously, to which the respective product mixtures are also continuously fed. During the processing of the liquid phase containing the respective target product and / or during the cooling of the respective product gas mixtures from the target product and the low boiling solvent, the separated high boiling solvents are removed. The boiling solvent is also conveniently continuously fed from the separation zone. The diacrylic acid and polyacrylic acid (both Michael adducts) present in the high boiling solvent are subsequently decomposed by exposure to high temperatures, and the acrylic acid monomers produced therefrom can be returned to the separation zone.

[0146] Those constituents of the product gas mixture continuously transferred to the separation zone, which as pure compounds have a boiling point below the boiling point of pure water at an operating pressure of 1 bar, remain mainly in the gas phase in the separation zone (in the case of water itself and of constituents which under corresponding conditions have a boiling point higher than the boiling point of pure water, the specific operating conditions applied to transfer the target product from the respective product gas mixture to the liquid phase determine the respective amounts remaining in the residual gas mixture) and are contained in the residual gas mixture R, the total mass flow rate M of which is continuously led off from the separation zone. In particular, those compounds contained as pure compounds in the product gas mixture continuously transferred to the separation zone, which have a boiling point below 250 K at an operating pressure of 1 bar (such compounds include above all the respective components of the applied inert diluent gas), are the major part of the residual gas mixture R. At least 95 mol-% of the total molar amount of compounds which are components of the product gas mixture continuously transferred to the separation zone and which, as pure compounds, have a boiling point below 250 K at an operating pressure of 1 bar are ultimately contained, under the applicable transfer conditions (steady operating conditions), in the total mass flow M of the residual gas mixture R which is continuously led off from the separation zone.

[0147] In practice, the absorptive and / or condensable conversion of the respective target product from the gas phase to the liquid phase is usually carried out in a separation column containing separation internals (to increase the mass transfer surface area). Useful separation internals include all known internals. In other words, trays, such as bubble cap trays, dual flow trays or valve trays, random packings, such as Raschig rings, or structured packings, such as Sulzer packings, can be used as separation internals. The respective product gas mixture is usually led upwards from below into the separation column. In the situation of absorptive condensation, the absorbent usually moves (is led) from above into the separation column. The downward flowing liquid absorbent essentially forms a liquid phase containing the target product (and secondary components with higher and similar boiling points). This liquid phase is continuously withdrawn from the lower part of the absorption column by side withdrawal (preferred) and / or withdrawal from the sump. In fractional condensation, the relatively high boiling constituents of the product gas mixture (e.g., including the target product acrylic acid) condense and rise to it. The enriched target product, particularly the acrylic acid-containing condensate, is typically conducted away from the condensation column via a side-draw removal, essentially forming a liquid phase containing the target product (and secondary components having higher and similar boiling points). It should be appreciated that absorption and condensation can also be utilized on top of each other. This applies, for example, whenever heat is further extracted from the system in the absorption process by direct and / or indirect cooling.

[0148] The respective product gas mixture is preferably led to a separation column, the temperature of which is reduced by indirect cooling or by direct cooling or by direct and indirect cooling. Advantageously, indirect cooling is carried out in a manner known per se in an indirect heat exchanger, whereas direct cooling is usually carried out, for example, by spraying a pre-cooled absorbent or a pre-cooled bottom liquid from the separation column onto the respective product gas mixture in a quenching device. A common feature of the above-described absorption and / or condensation processes (separation processes) is that at the top of the particular separation column, including separation internals (into the lower section of which the respective product gas mixture is usually led, preferably after its preceding direct and / or indirect cooling as described), a residual gas stream remains, which mainly comprises those constituents of the product gas mixture which have a boiling point of <250 K at standard pressure (1 bar) (i.e. constituents which are difficult to condense or otherwise volatilize).

[0149] These remaining gas streams include, for example, molecular nitrogen, which is further used as an inert diluent gas in the partial oxidation, the remaining molecular oxygen in excess in relation to the reaction stoichiometry in the partial oxidation, the carbon oxides (CO and / or CO2) formed as by-products or further used as inert diluent gases in the starting reaction gas mixture 1, but also propene that was not converted in the partial oxidation, and optionally co-used diluent gases, such as methane, natural gas, ethane, ethene and propane. In the preferred case, this residual gas may already be a residual gas mixture R. However, said remaining residual gas stream may still contain significant amounts of water vapor. This H2O content can be easily reduced, if desired, by applying additional condensation measures (for example corresponding indirect cooling). In the literature, the resulting condensate is usually named "acid water" (which usually contains some of the target product acrylic acid and / or carboxylic acid by-products, such as acetic acid). In the above-mentioned cases, the gas stream remaining after such additional condensation is therefore a residual gas mixture R.

[0150] The composition C of the resulting residual gas mixture R as described is usually 0.01 to 1% by volume of propene, 1 to 10% by volume of molecular oxygen, 0.5 to 25% by volume of water, 0-20% by volume of molecular nitrogen, 0.5 to 40% by volume of carbon monoxide, 1 to 75% by volume of carbon dioxide, 0-92% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially propane only); and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid It is.

[0151] Advantageously, not the entire flow of the residual gas mixture R is led away from the separation zone, since, for example, a portion of the residual gas mixture R can be advantageously used to strip the target product (e.g. acrylic acid) from a liquid phase containing the target product (e.g. acrylic acid), e.g. acid water, or can be generated during further purification of the absorbent or condensate containing the target product, which is continuously withdrawn from the separation column. The target product-containing stripping gas is then preferably returned to the separation column again.

[0152] However, from the total mass flow rate M of the residual gas mixture R, which is continuously withdrawn from the separation zone, a partial stream (=gas stream) having the same composition as the residual gas mixture R enters as a feed stream to the first reaction stage to be continuously recycled for preparing the stream of the starting reaction gas mixture 1, and the difference between the total mass flow rate M of the residual gas mixture R, which is continuously withdrawn from the separation zone, and the mass flow rate of the recycled gas stream P is continuously discharged from the production process as an off-gas stream O, all of the above subject to the following: a) the mass flow rate of the recycle gas stream P is at least 65% but not more than 99.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; and b) The mass flow rate of the off-gas stream O is less than 35% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone.

[0153] For the process of the present invention, applicable absorption and / or condensation processes for transferring the target product acrolein, respectively the target product acrylic acid, from a similar product gas mixture into the liquid phase are described in detail, for example, in documents US-A 3,433,840, WO 2014 / 22044, WO 2003 / 068721 and DE 10213998 (all directed to the absorption of acrolein by aqueous absorbents), WO 2008 / 090190 and WO 2013 / 139590 (both directed to the fractional condensation of acrylic acid), WO 2011 / 000808 and DE-A 10336386 (both directed to the absorption of acrylic acid via organic solvents with high boiling points) and DE-A 3,429,391, US-A 7,183,428 and US-A 5,315,037 (all directed to the absorption of acrylic acid by aqueous absorbents).

[0154] Further removal of the respective target product from the liquid phase to the desired degree of purity of the target product can be achieved by a wide variety of different combinations of a wide variety of different thermal separation processes, which may for example be combinations of extraction, desiccation, crystallization, rectification, azeotropic distillation, azeotropic rectification, distillation and / or stripping processes, all these processes being part of the separation zone and advantageously coupled to the separation column via a recycle stream to the appropriate said column.

[0155] It should be recognized that all process steps carried out in the separation zone are carried out with the inhibition of the undesired polymerization of the respective target product. This procedure may be as described in the cited prior art. Among the totality of available acrylic acid process stabilizers, prominent positions are taken by dibenzo-1,4-thiazine (PTZ=phenothiazine), 4-hydroxy-2,2,6,6-tetramethylpiperidine-1-oxyl (4-OH-TEMPO) and p-methoxyphenol (MEHQ). These can be part of the liquid phase containing acrylic acid, for example, either alone or in pairs or as a three-substance mixture. Hydroquinone is an excellent inhibitor for acrolein. Usually, the total amount of polymerization inhibitors in each liquid phase is 0.001-2% by weight relative to the target product present therein. Molecular oxygen dissolved in the relevant liquid phase usually further helps to promote polymerization inhibition. [Brief description of the drawings]

[0156] [Figure 1]In this document, a schematic block diagram of the production method of the present invention for the production of acrylic acid as a target product is shown. The reference numbers in Figure 1 are defined as follows: 1: optional steam stream 2: optional one or more hydrocarbon streams selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane and ethane 3: propene feed stream 4: molecular oxygen source stream 5: starting reaction gas mixture 16: first reaction stage 7: product gas mixture 18: optional secondary gas stream 9: starting reaction gas mixture 210: second reaction stage 11: product gas mixture 212: separation zone 13: a stream of the target product acrylic acid of its desired purity (produced via post-treatment of the liquid phase 15 in the separation zone), for example fed to a storage tank 14: Off-gas stream O15 discharged from the production process and fed to the synthesis gas production: liquid phase, into which the target product acrylic acid passes from the product gas mixture 2 16: residue gas mixture R 17: recycle gas stream P18: optional flow 17+14 of the residue gas mixture R circulating in the separation zone: total mass flow rate of the residue gas mixture R continuously led off from the separation zone M19: high boiling solvents brought from the separation zone (for example separated from the target product acrylic acid via work-up of the liquid phase 15 in the separation zone) DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0157] Thus, the present application more specifically encompasses the following embodiments: 1. A process for the continuous production of either acrolein or acrylic acid as a target product from propene under steady-state operating conditions, comprising: A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 20% by volume (preferably 0 to 10%) of water, 0-20% by volume of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, and molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5. and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 30% by volume of water, 0 to 20% by volume (preferably 0 to 10%) of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio of CO2:CO ranging from 1 to 4, and - Molecular oxygen and molecular nitrogen in a molar ratio of O2:N2 of at least 0.25 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. Including, a) the mass flow rate of the recycle gas stream P is at least 65% but not more than 99.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 35% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.01 to 1% by volume of propene, 1 to 10% by volume of molecular oxygen, 0.5 to 25% by volume of water, 0 to 20% by volume (preferably 0 to 10%) of molecular nitrogen, 0.5 to 40% by volume of carbon monoxide, 1 to 75% by volume of carbon dioxide, 0-92% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane, and ethene; and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid That is, the method.

[0158] 2. The method of embodiment 1, wherein the reaction temperature in the first reaction stage is 280° C. to 480° C.

[0159] 3. The method of embodiment 1, wherein the reaction temperature in the first reaction stage is 290°C to 470°C.

[0160] 4. The method of embodiment 1, wherein the reaction temperature in the first reaction stage is 300° C. to 450° C.

[0161] 5. The method of any of the preceding claims, wherein the operating pressure in the first reaction stage is from 1 to 5 bar.

[0162] 6. The method according to any of the preceding claims, wherein the operating pressure in the first reaction stage is from 1 to 4 bar.

[0163] 7. The method of any of the preceding claims, wherein the operating pressure in the first reaction stage is from 1 to 3.5 bar.

[0164] 8. The process of any of the preceding embodiments, wherein the propene conversion in a single pass in the first reaction stage is in the range of 85 mol% to 99 mol%.

[0165] 9. The process of any one of the preceding embodiments, wherein the propene conversion in a single pass in the first reaction stage is in the range of 90 mol% to 99 mol%.

[0166] 10. The process of any one of the preceding embodiments, wherein the propene conversion in a single pass in the first reaction stage is in the range of 92 mol % to 98.5 mol %.

[0167] 11. The associated selectivity S of the formation of the target product acrolein in the first reaction step A The method according to any one of the preceding embodiments, wherein the amount of the cation exchange resin is 75 mol % to 99 mol %.

[0168] 12. The associated selectivity S of the formation of the target product acrolein in the first reaction step A The method according to any one of the preceding embodiments, wherein the amount of the cation exchange resin is 80 mol % to 99 mol %.

[0169] 13. The method of any one of the preceding claims, wherein the reaction temperature in the second reaction stage is between 210°C and 430°C.

[0170] 14. The method of any one of the preceding claims, wherein the reaction temperature in the second reaction stage is between 220°C and 410°C.

[0171] 15. The method of any one of the preceding claims, wherein the reaction temperature in the second reaction stage is between 230°C and 390°C.

[0172] 16. The method of any of the preceding claims, wherein the operating pressure in the second reaction stage is from 1 to 5 bar.

[0173] 17. The method of any of the preceding claims, wherein the operating pressure in the second reaction stage is 1 to 4 bar.

[0174] 18. The method of any of the preceding claims, wherein the operating pressure in the second reaction stage is from 1 to 3.5 bar.

[0175] 19. The process of any of the preceding embodiments, wherein the acrolein conversion in a single pass in the second reaction stage is in the range of 97 mol % to 99.95 mol %.

[0176] 20. The process according to any of the preceding embodiments, wherein the acrolein conversion in a single pass of the second reaction stage of the process of the invention is in the range of 98 mol% to 99.9 mol%.

[0177] 21. The process according to any of the preceding embodiments, wherein the acrolein conversion in a single pass of the second reaction stage of the process of the invention is in the range of 99 mol% to 99.8 mol%.

[0178] 22. The associated selectivity S for the formation of the target product acrylic acid, based on the propene converted, evaluated over both reaction steps. AA 22. The method of any one of the preceding claims, wherein the is 83 mol% to 99 mol%.

[0179] 23. The associated selectivity S for the formation of the target product acrylic acid, based on the propene converted, evaluated over both reaction steps. AA 22. The method of any one of the preceding claims, wherein the is 85 mol % to 99 mol %.

[0180] 24. The associated selectivity S for the formation of the target product acrylic acid, based on the propene converted, evaluated over both reaction steps. AA 22. The method of any one of the preceding embodiments, wherein the is 87 mol % to 99 mol %.

[0181] 25. The method according to any one of the preceding claims, wherein the following conditions are simultaneously met: a) the initial reaction gas mixture 1 contains 3-10% by volume of propene, 3.9-23% by volume of molecular oxygen, and contains molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.3-2.1, and (only if acrylic acid is the target product, otherwise b) does not apply); b) The initial reaction gas mixture 2 contains 3-10% by volume acrolein, 1.8-25% by volume molecular oxygen, and a molar ratio of O2:acrolein in the range of 0.6-2.5.

[0182] 26. The method according to any one of embodiments 1 to 24, wherein the following conditions are simultaneously met: a) the initial reaction gas mixture 1 contains 4-9% by volume of propene, 5.6-19.8% by volume of molecular oxygen, and contains molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.4-2.2, and (only if acrylic acid is the target product, otherwise b) is not applicable); b) The initial reaction gas mixture 2 contains 4-9% by volume acrolein, 2.8-18% by volume molecular oxygen, and a molar ratio of O2:acrolein in the range of 0.7-2.0.

[0183] 27. The method according to any one of embodiments 1 to 24, wherein the following conditions are simultaneously met: a) the initial reaction gas mixture 1 contains 5-8% by volume of propene, 7.5-16.8% by volume of molecular oxygen, and contains molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.5-2.1, and (only if acrylic acid is the target product, otherwise b) is not applicable); b) The initial reaction gas mixture 2 contains 4.5-8% by volume acrolein, 3.6-12% by volume molecular oxygen, and a molar ratio of O2:acrolein in the range of 0.8-1.5.

[0184] 28. The first fixed catalyst bed of the first reaction stage comprises a catalyst body, the active material of which is represented by general formula I Mo 12 Bi a Fe b X 1 c X 2 d X 3 e X 4 f O n (I) wherein the variables are defined as follows: X 1 = nickel and / or cobalt, X 2 = thallium, alkali metals and / or alkaline earth metals, X 3 = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten, X 4 = silicon, aluminium, titanium and / or zirconium, a=0.5~5, b = 0.01 to 5, c=0~10, d=0~2, e=0~8, f=0~10 and n = a number determined by the valence and frequency of the elements other than oxygen in I. 28. The method of any one of the preceding claims, wherein the at least one multimetal oxide is

[0185] 29. The first fixed catalyst bed of the first reaction stage comprises a catalyst body, the active material of which is represented by general formula I Mo 12 Bi a Fe b X 1 c X 2 d X 3 e X 4 f O n (I) wherein the variables are defined as follows: X 1 = nickel and / or cobalt, X 2 = thallium, alkali metals and / or alkaline earth metals, X 3 = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten, X 4 = silicon, aluminium, titanium and / or zirconium, a=0.5~5, b=2~4, c=3~10, d=0.02~2, e=0~5, f=0~10 and n = a number determined by the valence and frequency of the elements other than oxygen in I. 28. The method of any one of the preceding claims, wherein the at least one multimetal oxide is

[0186] 30. The first fixed catalyst bed of the first reaction stage comprises a catalyst body, the active material of which is represented by general formula II [Bi a" Z 2 b" O x" ] p" [Z 2 12 Z 3 c" Z 4 d" Fe e" Z 5 f" Z6 g" Z 7 h" O y" ] q" (II) wherein the variables are defined as follows: Z 2 = molybdenum or molybdenum and tungsten, Z 3 = nickel and / or cobalt, Z 4 = thallium, alkali metals and / or alkaline earth metals, Z 5 = phosphorus, arsenic, boron, antimony, tin, cerium and / or lead, Z 6 = silicon, aluminium, titanium and / or zirconium, Z 7 = copper, silver and / or gold, a"=0.1~1, b"=0.2~2, c"=3~10, d"=0.02~2, e"=0.01~5, f"=0~5, g"=0~10, h"=0~1, x", y" = a number determined by the valence and frequency of elements other than oxygen in II, p" and q" are numbers whose p" / q" ratio is between 0.1 and 5. and at least one multimetal oxide of Bi, whose composition is different from these local environments and therefore separated from these local environments, whose maximum diameter (the longest direct line passing through the center of the domain and joining two points on the surface (interface) of the domain) is between 1 nm and 100 μm. a" Z 2 b" Ox " 28. The method of any one of the preceding claims, wherein the metal oxide is a multi-metal oxide comprising three-dimensional domains of:

[0187] 31. In the general formula II, Z 2 b" = (tungsten)b'' and Z 2 12 = (Molybdenum) 12 31. The method of embodiment 30, wherein

[0188] 32. The second fixed catalyst bed of the second reaction stage comprises a catalyst body, the active material of which is represented by general formula III Mo 12 V a X 1 b X 2 c X 3 d X 4 e X 5 f X 6 g O n (III) wherein the variables are 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=1~6, b=0.2~4, c=0.5~18, d=0~40, e=0~2, f=0~4, g=0~40 and 32. The method of any one of the preceding claims, wherein n=a number determined by the valence and frequency of the elements other than oxygen in III.

[0189] 33. The method of embodiment 32, wherein the variables of general formula III are each defined as follows: 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=1.5~5, b=0.5~2, c=0.5~3, d=0~2, e=0~0.2, f=0~1, g=0~40 and n = a number determined by the valence and frequency of the elements other than oxygen in III.

[0190] 34. The second fixed catalyst bed of the second reaction stage comprises a catalyst body, the active material of which is Mo 12 V a' Y 1 b' Y 2 c' Y 5 f' Y 6 g' O n' (IV) wherein the variables are defined as follows: 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'=2~4, b'=1~1.5, c'=1~3, f'=0~0.5, g'=0~8 and n' = a number determined by the valence and frequency of elements other than oxygen in IV. 32. The method of any one of the preceding claims, wherein the at least one multimetal oxide is of general formula IV:

[0191] 35. The process of any of the preceding embodiments, wherein the hourly space velocity of propene over the first fixed catalyst bed in the first reaction stage is from 50 to 300 l(STP) / l·h.

[0192] 36. The process of any of the preceding embodiments, wherein the hourly space velocity of propene over the first fixed catalyst bed in the first reaction stage is from 90 to 200 l(STP) / l·h.

[0193] 37. The process of any of the preceding embodiments, wherein the hourly space velocity of propene over the first fixed catalyst bed in the first reaction stage is from 135 to 200 l(STP) / l·h.

[0194] 38. The process of any of the preceding embodiments, wherein the hourly space velocity of acrolein over the second fixed catalyst bed in the second reaction stage is 40 to 290 l(STP) / l·h.

[0195] 39. The process of any of the preceding embodiments, wherein the hourly space velocity of acrolein over the second fixed catalyst bed in the second reaction stage is 80 to 190 l(STP) / l·h.

[0196] 40. The process of any of the preceding embodiments, wherein the hourly space velocity of acrolein over the second fixed catalyst bed in the second reaction stage is 125 to 190 l(STP) / l·h.

[0197] 41. The process of any of the preceding embodiments, wherein the shaped catalyst bodies of the first fixed catalyst bed in the first reaction stage are unsupported rings and the second fixed catalyst bed is a coated ring.

[0198] 42. The method of embodiment 41, wherein the unsupported ring has an outer diameter of 2 to 10 mm, a length of 2 to 10 mm, and a wall thickness of 1 to 3 mm.

[0199] 43. The method of embodiment 41 or 42, wherein the coated ring has an outer diameter of 4 to 10 mm, a length of 2 to 10 mm, and a wall thickness of 1 to 4 mm.

[0200] 44. Starting reactant gas mixture 1 is - recycled gas flow P; - Propene source; - A source of molecular oxygen; - optionally a flow of water vapor; and - optionally one or more streams of one or more hydrocarbons selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane and ethene; 44. The method of any one of the preceding claims, wherein the feed stream comprises

[0201] 45. The method of any of the preceding embodiments, wherein the initial reactant gas mixture 1 is formed from a feed stream containing oxygen as a source of molecular oxygen with a purity of at least 90% by volume O2 and no more than 10% by volume N2.

[0202] 46. ​​The method of any of the preceding embodiments, wherein the initial reactant gas mixture 1 is formed from a feed stream containing oxygen as the source of molecular oxygen with a purity of at least 95% by volume O2 and no more than 5% by volume N2.

[0203] 47. The method of any of the preceding embodiments, wherein the initial reactant gas mixture 1 is formed from a feed stream containing oxygen as a source of molecular oxygen with a purity of at least 98% by volume O2 and no more than 2% by volume N2.

[0204] 48. The method of any of the preceding embodiments, wherein the initial reactant gas mixture 1 is formed from a feed stream containing oxygen as a source of molecular oxygen with a purity of at least 99% by volume O2 and 1% by volume or less N2.

[0205] 49. The method of any of the preceding claims, wherein the initial reaction gas mixture 1 is formed from a feed stream containing oxygen as a source of molecular oxygen with a purity of 99.4% by volume to 99.9% by volume O2 and 0.1% by volume to 0.6% by volume N2.

[0206] 50. The method of any of the preceding claims, wherein the initial reactant gas mixture 1 is formed from a feed stream containing oxygen with a purity of 99.5% by volume to 99.8% by volume O2 and 0.2% by volume to 0.5% by volume N2 as a source of molecular oxygen.

[0207] 51. The method according to any of the preceding claims, wherein as a secondary gas, oxygen with a purity of at least 90% by volume O2 and no more than 10% by volume N2 is added to the product gas mixture 1 leaving the first reaction stage.

[0208] 52. The method according to any of the preceding claims, wherein as a secondary gas, oxygen with a purity of at least 95% by volume O2 and not more than 5% by volume N2 is added to the product gas mixture 1 leaving the first reaction stage.

[0209] 53. The method according to any of the preceding claims, wherein as a secondary gas, oxygen with a purity of at least 98% by volume O2 and not more than 2% by volume N2 is added to the product gas mixture 1 leaving the first reaction stage.

[0210] 54. The method according to any of the preceding embodiments, wherein as a secondary gas, oxygen with a purity of at least 99% by volume O2 and not more than 1% by volume N2 is added to the product gas mixture 1 leaving the first reaction stage.

[0211] 55. The method according to any of the preceding claims, wherein as secondary gas, oxygen with a purity of 99.4% by volume to 99.9% by volume O2 and 0.1% by volume to 0.6% by volume N2 is added to the product gas mixture 1 leaving the first reaction stage.

[0212] 56. The method according to any of the preceding claims, wherein as secondary gas, oxygen with a purity of 99.5% by volume to 99.8% by volume O2 and 0.2% by volume to 0.5% by volume N2 is added to the product gas mixture 1 leaving the first reaction stage.

[0213] 57. The method of any of the preceding embodiments, wherein the initial reactant gas mixture 1 is formed from a feed stream comprising a propene feedstock containing ≧99 vol.% propene and ≦1 vol.% propane.

[0214] 58. The method of any of the preceding embodiments, wherein the initial reaction gas mixture 1 is formed from a feed stream comprising a propene feedstock containing ≧90 vol.% propene and ≦10 vol.% but ≧2 vol.% propane.

[0215] 59. The method of any of the preceding embodiments, wherein the initial reaction gas mixture 1 is formed from a feed stream comprising a propene feedstock containing ≧70 vol.% propene and ≦30 vol.% but ≧15 vol.% propane.

[0216] 60. The method of any of the preceding embodiments, wherein an inert gas is added as a secondary gas to the product gas leaving the first reaction stage.

[0217] 61. The method of any of the preceding embodiments, wherein the initial reactant gas mixture 1 is formed from a feed stream comprising one or more streams of at least one hydrocarbon selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane, and ethene.

[0218] 62. The method according to any of the preceding claims, wherein the mass flow rate of the recycle gas stream P is at least 70% but not more than 99.5% of the mass flow rate M of the residual gas mixture R continuously led off from the separation zone, and the mass flow rate of the off-gas stream O is not more than 30% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone.

[0219] 63. The method according to any of the preceding claims, wherein the mass flow rate of the recycle gas stream P is at least 80% but not more than 99.0% of the mass flow rate M of the residual gas mixture R continuously led from the separation zone, and the mass flow rate of the off-gas stream O is not more than 20% but at least 1.0% of the total mass flow rate M of the residual gas mixture R continuously led from the separation zone.

[0220] 64. The method of any of the preceding claims, wherein the mass flow rate of the recycle gas stream P is at least 95% but not more than 98.0% of the mass flow rate M of the residual gas mixture R continuously led from the separation zone, and the mass flow rate of the off-gas stream O is not more than 5% but at least 2.0% of the total mass flow rate M of the residual gas mixture R continuously led from the separation zone.

[0221] 65. Under steady-state operating conditions, A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 15 to 25% by volume of carbon monoxide, 20 to 55% by volume of carbon dioxide, 10-30% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially only propane); and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, - and molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5 and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 20% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 15 to 25% by volume of carbon monoxide, 20 to 55% by volume of carbon dioxide, 10-30% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially only propane); and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio of CO2:CO ranging from 1 to 4, and - Molecular oxygen and molecular nitrogen in a molar ratio of O2:N2 of at least 0.25 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led from the separation zone to the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing a gas flow P, as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R, a) the mass flow rate of the recycle gas stream P is at least 90% but not more than 99% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 10% but at least 1% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.05 to 0.5% by volume of propene, 1-4% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 15 to 30% by volume of carbon monoxide, 30-60% by volume of carbon dioxide, 10-50% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially only propane); and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid 62. The method of any preceding embodiment, wherein

[0222] 66. The method of embodiment 65, wherein the initial reaction gas mixture 1 is formed from a feed stream consisting of a propene feedstock stream containing ≧90 vol.% propene and ≦10 vol.% but ≧2 vol.% propane, a molecular oxygen stream with a purity of ≧99.0 vol.% and ≦99.9 vol.% O2 and ≦1 vol.% and ≧0.1 vol.% N2, and a recycle gas stream P.

[0223] 67. Under steady-state operating conditions, A) As an essential step, in the first reaction step, 2 to 15 volume percent propene, 2.4 to 37.5% by volume of molecular oxygen; 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 1.2 to 15 volume percent carbon monoxide, 3 to 30% by volume of carbon dioxide, 40-80% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol % of the total molar amount of these hydrocarbons is methane), and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - molecular oxygen and propene in a molar ratio of O2:propene in the range of 1.2 to 2.5, - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, - molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5, and - Carbon monoxide and total molar amounts of hydrocarbons in the CO:total hydrocarbons molar ratio range of >0-0.3 and a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and a concomitant selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe) such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2 to 15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2 to 20% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 1.2 to 15% by volume of carbon monoxide, 3 to 30% by volume of carbon dioxide, 40-80% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol % of the total molar amount of these hydrocarbons is methane), and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - molecular oxygen and acrolein in a molar ratio of O2:acrolein ranging from 0.5 to 3; - carbon dioxide and carbon monoxide in a molar ratio CO2:CO ranging from 1 to 4, - molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.25, and - Carbon monoxide and total molar amounts of hydrocarbons (including acrolein) in the CO:total hydrocarbons molar ratio range of >0-0.3 The thus-produced initial reaction gas mixture 2 stream, containing AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing a gas stream P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas stream P from the production process as an off-gas stream O having the same composition as the residual gas mixture R, Including, a) the mass flow rate of the recycle gas stream P is at least 75% but not more than 98% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 25% but at least 2% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.05 to 0.8 volume percent propene, 1-4% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0 to 10% by volume (preferably 0 to 5%) of molecular nitrogen, 1.2 to 15% by volume of carbon monoxide, 3 to 30% by volume of carbon dioxide, 45-90% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably at least 80 mol % of the total molar amount of these hydrocarbons is methane), and 0.01-5% by volume of other compounds, including but not limited to acrolein and / or acrylic acid 62. The method of any preceding embodiment, wherein

[0224] 68. The method according to embodiment 67, wherein the initial reaction gas mixture 1 is formed from a feed stream consisting of a polymer grade (propene feedstock containing ≧99% by volume of propene and ≦1% by volume of propane) or chemical grade propene (propene feedstock containing ≧90% by volume of propene and ≦10% by volume but ≧2% by volume of propane), a molecular oxygen stream with a purity of ≧99.0% by volume and ≦99.9% by volume of O2 and ≦1% by volume and ≧0.1% by volume of N2, a recycle gas stream P, and an additional stream of at least one hydrocarbon selected from the group consisting of methane, natural gas, ethene, propane, n-butane, isobutane and ethene.

[0225] 69. The method of embodiment 68, wherein at least 80 mol % of the total molar amount (amount) of hydrocarbons in the at least one hydrocarbon additional stream is methane or natural gas.

[0226] 70. The method of embodiment 68, wherein at least 90 mol % of the total molar amount (amount) of hydrocarbons in the at least one hydrocarbon additional stream is methane or natural gas.

[0227] 71. The method of embodiment 68, wherein at least 95 mol % of the total molar amount (amount) of hydrocarbons in the at least one hydrocarbon additional stream is methane or natural gas.

[0228] 72. The method of embodiment 68, wherein 100 mol % of the total molar amount (amount) of hydrocarbons in the at least one hydrocarbon additional stream is methane or natural gas.

[0229] 73. The method of any of the preceding embodiments, wherein starting reactant gas mixture 1 and starting reactant gas mixture 2 are non-explosive.

[0230] 74. The ratio Q in the starting reactant gas mixture 1 = (n CO ) / (n CO +n 炭化水素 74. The method of any one of the preceding claims, wherein:

[0231] 75. The ratio Q in the starting reactant gas mixture 1 = (n CO ) / (n CO +n 炭化水素 74. The method of any one of the preceding claims, wherein the ratio of the saturation peak to the peak width is in the range of 0.005 to 0.7.

[0232] 76. The ratio Q in the starting reactant gas mixture 1 = (n CO ) / (n CO +n 炭化水素 74. The method of any one of the preceding claims, wherein the ratio of the number of saturation saturation regions to the number of saturation regions is in the range of 0.005 to 0.6.

[0233] 77. The ratio Q in the starting reactant gas mixture 2 = (n CO ) / (n CO +n 炭化水素 77. The method of any one of the preceding claims, wherein the ratio of 1 to 76 is in the range of >0 to 0.8.

[0234] 78. The ratio Q in the starting reactant gas mixture 2 = (n CO ) / (n CO +n 炭化水素 77. The method of any one of the preceding claims, wherein the ratio of the number of saturation saturation regions to the number of saturation regions is in the range of 0.005 to 0.7.

[0235] 79. The ratio Q in the starting reactant gas mixture 2 = (n CO ) / (n CO +n 炭化水素 77. The method of any one of the preceding claims, wherein the ratio of the saturation peak to the peak width is in the range of 0.005 to 0.6.

[0236] 80. The method of any of the preceding embodiments, wherein at least 96 mol % of each target product contained in each product gas mixture transferred to the separation zone is transferred to the liquid phase in the separation zone.

[0237] 81. The method according to any of the preceding embodiments, wherein at least 97 mol % of each target product contained in each product gas mixture transferred to the separation zone is transferred to the liquid phase in the separation zone.

[0238] 82. The method according to any of the preceding embodiments, wherein at least 98 mol % of each target product contained in each product gas mixture transferred to the separation zone is transferred to the liquid phase in the separation zone.

[0239] 83. The method according to any of the preceding embodiments, wherein at least 99 mol % of each target product contained in each product gas mixture transferred to the separation zone is transferred to the liquid phase in the separation zone.

[0240] 84. The method according to any of the preceding embodiments, wherein at least 99.5 mol % but not more than 99.8 mol % of each target product contained in each product gas mixture transferred to the separation zone is transferred to the liquid phase in the separation zone.

[0241] 85. The method according to any of the preceding claims, wherein the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone contains, as pure compounds, at least 96 mol % of the total molar amount of compounds contained in the product gas mixture continuously transferred to the separation zone, which have a boiling point below 250 K at an operating pressure of 1 bar.

[0242] 86. The method according to any of the preceding claims, wherein the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone contains, as pure compounds, at least 97 mol % of the total molar amount of compounds contained in the product gas mixture continuously transferred to the separation zone, which have a boiling point below 250 K at an operating pressure of 1 bar.

[0243] 87. The method according to any of the preceding claims, wherein the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone contains, as pure compounds, at least 98 mol % of the total molar amount of compounds contained in the product gas mixture continuously transferred to the separation zone, which have a boiling point below 250 K at an operating pressure of 1 bar.

[0244] 88. The method according to any of the preceding claims, wherein the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone contains, as pure compounds, at least 99 mol %, or at least 99.5 mol %, but less than 100 mol %, of the total molar amount of compounds contained in the product gas mixture continuously transferred to the separation zone that have a boiling point below 250 K at an operating pressure of 1 bar.

[0245] 89. The method of any of the preceding claims, wherein transferring each target product from the gas phase to the liquid phase comprises absorbing each target product into a liquid absorbent and / or condensing (preferably fractionally condensing) each target product from the gas phase in at least one separation column containing separating internals through which the respective product gas mixture is directed upwardly from below.

[0246] 90. The method of any of the preceding claims, wherein the liquid phase containing the respective target product in the separation zone is subjected to a work-up process including separation of the high boiling solvent from the target product, and the high boiling solvent and purified target product are continuously withdrawn from the separation zone to produce the target product in its desired purity.

[0247] 91. A production method according to any one of the preceding embodiments, characterized in that it operates in combination with another process and comprises a step of converting an off-gas stream O discharged from the production process and fed to said other process into a synthesis gas, and / or mixing said off-gas stream O with another synthesis gas produced in said other process from any feedstock other than the off-gas O, to produce the other synthesis gas thus modified.

[0248] 92. The conversion of the off-gas stream O2 to at least one of the following reactions: - heterogeneously catalyzed steam reforming in the gas phase of the hydrocarbons contained in the off-gas stream O; - heterogeneously catalyzed hydrogenation in the gas phase of the CO2 contained in the off-gas stream O; - heterogeneously catalyzed dry reforming in the gas phase of hydrocarbons with CO2, both of which are contained in an off-gas stream O, as well as - partial gas phase combustion of hydrocarbons and / or compounds composed of C, H and O and contained in the off-gas stream O 92. The method of producing according to embodiment 91, comprising:

[0249] 93. The method of production according to embodiment 91 or 92, wherein the synthesis gas produced and / or other modified synthesis gas is subsequently converted in a synthesis method to produce any chemical compound containing C and H, and optionally O.

[0250] Examples and Comparative Examples The production plant used for the following examples 1-4 and comparative examples is always the same. This production plant comprises a partial oxidation section and a separation zone. The design of the separation zone including its apparatus qualitatively corresponds to that of example 1 of WO 2008 / 090190, differing therefrom only in its dimensions. The partial oxidation section is essentially an oxidation reactor system comprising two fixed-bed two-zone multitubular (tube bundle) reactors in series, including an aftercooler connected downstream of the first stage reactor for indirect cooling of the product gas mixture 1 leaving the first stage multitubular reactor. The design of the reactor system including the salt melt acting as the heat exchange medium qualitatively corresponds to that described in US-A0249196, sections

[0187] to

[0226] . Between the outlet "aftercooler" and the inlet "second stage reactor" is a means for feeding compressed secondary molecular oxygen-containing gas. In Examples 2-4 and Comparative Example, elements that do not differ from the corresponding elements described in Example 1 are no longer repeated in Examples 2-4 and Comparative Example. The data presented for the separation zone in the illustrative examples / comparative examples are derived from computational simulations and not from actual experiments. The data presented for the partial oxidation section are derived from results obtained in comparable single tube experiments applying corresponding catalyst loadings and similar / equivalent starting reaction gas mixtures and reaction conditions. EXAMPLES

[0251] Describes steady state (small amounts of O2 dissolved in the liquid material stream are not included (not mentioned) below)

[0252] The packed contact tubes for the first reaction stage (propene → acrolein) in the two-zone tube bundle reactor correspond to the following reaction tubes: The reaction tube (V2A steel; outer diameter: 30 mm, wall thickness: 2 mm, inner diameter: 26 mm, length: 350 cm) is filled from top to bottom as follows:

[0253] Section 1: As a spare bed, a 50 cm long ring of steatite (C220 steatite, CeramTec GmbH, D-73207 Plochingen) with dimensions 7 mm x 3 mm x 4 mm (external diameter x length x internal diameter).

[0254] Section 2: A catalyst packing of 35 wt.-% steatite (C220 steatite, CeramTec GmbH, D-73207 Plochingen) rings, shape 5 mm x 5 mm x 2 mm (external diameter x length x internal diameter) and 115 cm long of a homogeneous mixture of 65 wt.-% of the unsupported catalyst of section 3.

[0255] Section 3: 185 cm long catalyst packing of annular unsupported catalyst (5 mm x 5 mm x 2 mm = outer diameter x length x inner diameter), the geometry of which is produced as the catalyst described in Example 3 of US-B 6,881,702 (stoichiometry: Mo 12 Co7Fe 2.94 Bi 0.6 S 1.59 K 0.08 O x The mass and lateral compressive strength (determined as defined in WO 2005 / 030393 and WO 2007 / 017431) of the uncalcined precursor ring were 189±4 mg and 23±3 N, respectively.

[0256] The first 175 cm from top to bottom is pumped in countercurrent, at temperature T A The second 175 cm is pumped in countercurrent, at temperature T B The salt bath B is thermostatically regulated by means of a salt bath B having a

[0257] The packed contact tubes in the two-zone tube bundle reactor for the second reaction stage (acrolein → acrylic acid) are represented by the following reaction tubes: The reaction tube (V2A steel; outer diameter: 30 mm, wall thickness: 2 mm, inner diameter: 26 mm, length: 350 cm) is filled from top to bottom as follows:

[0258] Section 1: As a preliminary bed, a 20 cm long ring of steatite (C220 steatite, CeramTec GmbH, D-73207 Plochingen) with the dimensions 7 mm x 3 mm x 4 mm (external diameter x length x internal diameter).

[0259] Section 2: 140 cm long catalyst packing of annular coated catalyst (7 mm x 3 mm x 4 mm = outer diameter x length x inner diameter), produced as catalyst of example C8 of US-A2015 / 0080605, with an oxidized eggshell content of 15 wt. % and an active material stoichiometry of Mo 12 V3W 1.2 O x It is.

[0260] Section 3: A catalyst packing of 190 cm length (7 mm x 3 m x 4 mm = outer diameter x length x inner diameter) of annular coated catalyst, produced as catalyst of example C3 of US-A 2015 / 0080605, with an oxidized eggshell content of 20 wt. % and an active material stoichiometry of Mo 12 V3W 1.2 O x It is.

[0261] The first 175 cm from top to bottom is pumped in countercurrent, at temperature T C The second 175 cm is pumped in countercurrent, at temperature T D The temperature is thermostatically regulated by means of a salt bath D having a

[0262] Starting reactant gas mixture 1 is a mixture of the following feed gas streams: - 15760 kg / h of molecular oxygen (purity: 99.9999% by mass O2); - 16000 kg / h of chemical grade propene (94.77% by weight propene and 5.23% by weight propane); and - a recycled gas flow P of 211617 kg / h.

[0263] The temperature of the initial reaction gas mixture 1 is 200° C. and its pressure (input pressure) is 2.434 bar. The composition of the initial reaction gas mixture 1 fed through the first stage tube bundle reactor of reaction tubes is as follows:

[0264] [Table 1]

[0265] temperature T A is 327°C. Temperature T B is 335°C.

[0266] The temperature of the catalytically active portion of the fixed catalyst bed in the reaction tube is in the range of 325°C to 398°C (the lower limit of the range indicates the lowest temperature of the catalytically active portion of the fixed catalyst bed, and the upper limit of the range indicates the highest temperature of the catalytically active portion of the fixed catalyst bed).

[0267] The hourly space velocity of propene over the fixed catalyst bed (section 2 + section 3 in the reactor tube) was 164 (l(STP) / l·h).

[0268] The single-pass propene conversion was 95.85 mol % with a concomitant selectivity for the formation of acrolein S A is 91.64 mol%.

[0269] The flow of 243377 kg / h of product gas mixture 1 leaving the first reaction stage has the following composition at a temperature of 360° C. and a pressure of 1.834 bar:

[0270] [Table 2]

[0271] The temperature of the product gas mixture leaving the first reaction stage (product gas mixture 1) is reduced by indirect cooling in an aftercooler. The thus cooled product gas mixture 1 stream (temperature = 250 °C) is mixed with a secondary molecular oxygen stream of 3612.9 kg / h (purity: 99.9999% by weight O2; temperature: 165.00 °C; pressure: 2.50 bar).

[0272] The composition of the thus produced starting reaction gas mixture 2 (temperature: 249.2° C.; pressure: 1.809 bar), fed through the reaction tubes in the second-stage tube bundle reactor, of 246990 kg / h is as follows:

[0273] [Table 3]

[0274] temperature T C is 267°C. Temperature T D is 272°C.

[0275] The temperature of the catalytically active portion of the fixed catalyst bed in the reaction tube is in the range of 265°C to 302°C (the lower limit of the range indicates the lowest temperature of the catalytically active portion of the fixed catalyst bed, and the upper limit of the range indicates the highest temperature of the catalytically active portion of the fixed catalyst bed).

[0276] The hourly space velocity of acrolein over the fixed catalyst bed (sections 2 and 3 of the reactor tube) is 132 (l(STP) / l·h).

[0277] The single pass acrolein conversion was 99.53 mol % and the associated selectivity for the formation of acrylic acid, S, based on the converted propene, was evaluated over both reaction stages. AA is 89.77 mol%.

[0278] Leaving the second stage reactor (product gas mixture 2) and at its inlet into the separation zone, the product gas mixture (246990 kg / h, T=250° C., P=1.509 bar) is cooled to a temperature of 120° C. by direct cooling in a spray cooler (quench 1) operating in parallel current.

[0279] The composition of product gas mixture 2 is as follows:

[0280] [Table 4]

[0281] The liquid to be used for direct cooling of the product gas mixture 2 (quench liquid 1) is a part (529660 kg / h, T=116.5° C.) of the bottom liquid withdrawn from the bottom of the condensation column described below in a total amount of 533070 kg / h (T=116.5° C.). The content of this bottom liquid is:

[0282] [Table 5]

[0283] The mixture of product gas mixture cooled to 120° C. and unevaporated quench liquid 1, which has the same temperature as a result of direct cooling, is thus led to the bottom of the condensation column. The pressure in the bottom space and in the quench 1 is 1.479 bar. The other part of the bottom liquid withdrawn from the bottom of the condensation column (3410 kg / h) is fed to the second stripping column and fed to its central section. The second stripping column contains dual flow trays as separating internals. Just like the condensation column, the second stripping column is thermally insulated from the environment.

[0284] The energy required for cracking the diacrylic acid and polyacrylic acid (both Michael adducts) contained in the bottom liquid to acrylic acid is replenished to the second stripping column by means of an external forced circulation three-flow tube bundle flash evaporator, which is fed with the bottom liquid of the second stripping column. The heat carrier conducted through the space around the heat exchanger tubes is pressurized with water vapor. As this heat carrier flows through the heat exchanger tubes, the bottom liquid is heated, the main amount of which is then recycled to the bottom of the second stripping column. A small part of the total amount of bottom liquid conducted through the heat exchanger is branched, degassed, diluted with methanol and sent to residue incineration.

[0285] In addition, the first gaseous content withdrawn from the top of the first stripping column is fed to the bottom of the second stripping column. The second gaseous content (particularly containing acrylic acid and water) is withdrawn from the top of the second stripping column in an amount of 18851 kg / h (temperature=65° C., pressure=1.40 bar) and fed to the bottom of the quench 1 and / or condensation column. The content of the second gaseous content is as follows:

[0286] [Table 6]

[0287] A centrifugal droplet separator is integrated into the bottom space of the condensation column, which prevents droplets of bottom liquid from being entrained upwards from the bottom space.

[0288] The bottom space of the condensation column is completed with a first collection tray (a stack tray with a uniformly distributed covered stack).

[0289] The collection tray has a double wall construction with an inward 2° slope and a central extraction cup and extraction nozzle. The free gas cross-sectional area is approximately 30%.

[0290] 104368 kg / h of the high boiler fraction is led from this first collecting tray into a bottom space arranged below the first collecting tray. The high boiler fraction has the following content at a temperature of 105.8° C. and a pressure of approximately 1.479 bar:

[0291] [Table 7]

[0292] The bottom temperature is 116.5° C. and the bottom pressure (liquid level) is 1.479 bar.

[0293] Above the first collecting tray, a first sequence of equidistant dual flow trays is arranged. The passage openings consist of annular openings of uniform diameter, with the punching burrs facing downwards in the separation column. The arrangement of the centers of the passage circles follows a strict triangular pitch. The top tray serves as a distributor tray. For this purpose, the wall of the column contains two inserted tubes with holes between the second collecting tray and the aforementioned top tray. The crude acrylic acid and the mother liquor are recycled to the condensation column via the inserted tubes. The first series of dual flow trays is completed with a second collecting tray (stack tray with uniformly distributed covered stack; central withdrawal cup with withdrawal nozzle, gas-free cross-sectional area about 30%), which is fixed above the last dual flow tray.

[0294] From this second collecting tray, 266895 kg / h of crude acrylic acid is continuously withdrawn as a first side draw at a temperature of 103.6° C. and 1.443 bar, which has the following content:

[0295] [Table 8]

[0296] The 60,745 kg / h of crude acrylic acid withdrawn from the second collecting tray is recycled to the condensation column via the respective inserted tubes immediately below the second collecting tray, together with the mother liquor (68,362 kg / h) obtained by further purification by crystallization of the withdrawn acrylic acid and heated to 90° C. by indirect heat exchange with the withdrawn crude acrylic acid and steam as a heat medium, to the dual flow tray following below the second collecting tray.

[0297] 115000 kg / hr of crude acrylic acid withdrawn from the second collection tray is recycled directly onto the second collection tray via spray nozzles at a temperature of 103.6° C. to keep this section wet and prevent contamination.

[0298] The 91,149 kg / h of crude acrylic acid withdrawn from the second collecting tray is cooled to a temperature of 29° C. by multi-stage indirect heat exchange (among other things in a thermally integrated manner for recycling the mother liquor to the condensation column mentioned above) and, optionally, intermediately buffered in a tank farm. About 1,000 kg / h of deionized water is then added to the cooled acrylic acid.

[0299] The resulting mixture is cooled and subsequently crystallized in a cooled disk crystallizer, and the resulting crystal suspension is purified in a hydraulic melt wash column as described on pages 49 and 50 of WO 2008 / 090190. From the melt circuit, stabilized by the addition of a solution of MEHQ in glacial acrylic acid, 22787 kg / h of the desired separated acrylic acid are withdrawn as glacial acrylic acid, which has the following contents:

[0300] [Table 9]

[0301] It is eminently suitable for preparing superabsorbents based on sodium poly-acrylate. Inhibitor solution 1 is prepared at 25° C. by dissolving 11 kg / h of PTZ (phenothiazine) in 904 kg / h of the abovementioned glacial acrylic acid. The remaining stream of glacial acrylic acid, which is continuously withdrawn from the melting circuit, is continuously fed to a storage tank.

[0302] The mother liquor removed in the wash column is first conducted to a heatable collecting vessel and from there to a tank from which it is heated to 90° C. by thermal integration (as already described) and recycled to the condensation column with a mass flow rate of 68,362 kg / h together with 60,745 kg / h of crude acrylic acid withdrawn on the second collecting tray of the condensation column (as a mixed stream (96.4° C., 1.10 bar)) in the upper part of the series of dual flow trays below the second collecting tray.

[0303] The composition of the recycled mother liquor is as follows:

[0304] [Table 10]

[0305] Above the second collecting tray of the condensation column, the first dual flow tray of the second arrangement of equidistant dual flow trays is placed. The passage openings consist of annular openings of uniform diameter, with the punching burrs facing downwards in the separation column. The arrangement of the centers of the passage circles follows a strict triangular pitch. Above the last dual flow tray, the equidistant arrangement of conventional single flow Thormann trays begins. The very first Thormann tray from the bottom is the tray from which the liquid leaving the tray flows out through downcomers configured as tubes. These tubes are hydraulically sealed from the gas space of the next dual flow tray below. The height of the drain weirs is reduced in the flow direction of the cross-flow tray. The hydraulic seal has empty openings due to the impingement plate. The drains are uniformly distributed in the last third of the tray cross section (opposite the feed on the tray). The hydraulic seal is made into a cup with an inclined overflow weir (45°). Alternatively, the Thorman tray may be configured such that opposing flow directions of liquid are achieved in successive channels through the arrangement of drive slots in the hood of the Thorman tray.

[0306] A third collection tray (a stack tray with approximately evenly distributed covered stack) is positioned above the upper Thorman tray. 843612 kg / hr of acid water is withdrawn as a second side draw from the third collection tray at a temperature of 60.7° C. and a pressure of 1.229 bar.

[0307] Acid water has the following content:

[0308] [Table 11]

[0309] The withdrawn 34405 kg / h of acid water (60.7° C.) are recycled to the top Soman tray together with 35 kg / h of inhibitor solution 1 (25° C.) and 28 kg / h of molten MEHQ (T=80° C.). 880 kg / h of inhibitor solution 1 are recycled (temperature 25° C.) to the single-flow Soman tray arrangement (after approximately two-thirds of the length of this Soman separation section (counting from its bottom)).

[0310] Pulled out 458.80m 3 / hour of acid water, at a temperature of 28.1°C, is recycled to the centre of an equidistant arrangement of valve trays, which are described in more detail below (cooling is effected by means of multi-stage indirect heat exchange).

[0311] The withdrawn 325000 kg / h of acid water are recycled at a temperature of 23.2° C. to the top of the above-mentioned equidistant arrangement of valve trays (cooling is carried out together with the above-mentioned amount of acid water by means of multi-stage indirect heat exchange; the final cooling stage from 28.1° C. to 23.2° C. is carried out thermally and with the aid of heat integration (liquid chemical-grade propene is used as coolant and is evaporated simultaneously; the gaseous propene produced is subsequently used to constitute the starting reaction gas mixture 1 for the first stage of the heterogeneously catalyzed gas-phase partial oxidation of propene).

[0312] The withdrawn 9207 kg / h of acid water is fed to an extraction column for the purpose of subsequent extraction.

[0313] Above the third collecting tray in the condensation column, a series of two-flow valve trays are fixed in an equidistant arrangement. The overflow weir height of the upper valve tray is higher than that of the lower valve tray. The valves used are VV12 valves from Stahl, Viernheim, Germany. The pressure at the top of the condensation column is 1.17 bar. At the top of the condensation column, 231380 kg / h of the residual gas mixture R leaves the separation column at a temperature of 24° C. and the following content:

[0314] [Table 12]

[0315] In the indirect heat exchanger, the gas mixture R is heated to 32° C. 3763 kg / h thereof are discharged as off-gas stream O from the production process.

[0316] 227617 kg / h of said gas mixture R are compressed by means of a compressor to a pressure of 2.5 bar, raising its temperature to approximately 150° C. 16000 kg / h of it are fed to a first stripping column in the separation zone for stripping the extract from the acid water extraction. The remaining 211617 kg / h of the compressed gas mixture (as cycle gas) are recycled as gas stream P to the first reaction stage 1 as feed stream for preparing the stream of starting reaction gas mixture 1. The sum of the mass flows of the off-gas stream O and the cycle gas stream P constitutes the total mass flow M of the remaining gas mixture R, which is continuously led off from the separation zone. This total mass flow M contains 99.99 mol % of the total molar amount of those compounds contained in the product gas mixture 2, which is continuously transferred to the separation zone, as pure compounds and which have a boiling point below 250 K at a working pressure of 1 bar.

[0317] The mass flow rate of the recycle gas stream P is 98.37% of the total mass flow rate M of the remaining gas mixture R, which is continuously led off from the separation zone.

[0318] The extraction column for the acid water extraction contains as separating internals punched structured packings made of stainless steel sheets (material 1.4571) of Montz-Pak B1-350 type, fitted flush with the edges, and arranged one on top of the other. The internal diameter of the extraction column is uniform throughout the packing. The extraction solvent used is Palatinol® A (diethyl phthalate). The internal diameter of the bottom and top vessel of the extraction column is enlarged (compared to the uniform internal diameter of the column along the packing) in order to improve the phase separation at the bottom and to reduce the entrainment of the extraction solvent at the top of the column. In addition, a random plastic bed (e.g. polyethylene or Teflon®) is introduced as a coalescence aid at the top of the column. 9207 kg / h of acid water (temperature = 60.7 °C) to be extracted is fed to the extraction column below the bottom packing via a tubular distributor with suitable passage openings. Above the top packing of the extraction column, a mixture of fresh Palatinol® A and the extraction solvent recycled from the first stripping column and previously stripped free therein (ratio of the respective mass flows 1:240) (temperature=50° C.) is introduced at a mass flow rate essentially equal to 9207 kg / h of acid water. The recycled extraction solvent has the following content:

[0319] [Table 13]

[0320] The specific mass of acid water (at temperature = 60.7°C) is 1011.5 kg / m 3The extraction solvent is introduced in the same way through a tubular distributor with suitable passage openings. The acid water forms a continuous phase, while the extraction solvent forms a dispersed phase in the form of droplets (liquid particle diameter is in the range of 2-5 mm), which descends within the aqueous phase. At the top of the extraction column, a raffinate is withdrawn. This raffinate is sent to incineration. The extract is withdrawn from the bottom of the extraction column. It contains the acrylic acid extracted from the acid water. The entire extract is led to the top of the first stripping column. The extract is preheated to 95 ° C by indirect heat exchange in a plate heat exchanger. The heat carrier is the bottom liquid withdrawn in the first stripping column. From the bottom to the top, the first stripping column contains as separating internals a first arrangement of equidistant dual flow trays, followed by an arrangement of three times larger single flow Thorman trays. Just like the extraction column, the first stripping column is insulated from the environment. The hole diameters of the dual flow trays are uniform (the arrangement of the holes corresponds to a strict triangular pitch). The Thorman trays are constructed in such a way that the opposing flow directions of the liquid are in each case obtained in successive channels in the cross-flow direction through the arrangement of drive slots in the hoods of the Thorman trays. Above the last Thorman tray there is also arranged a bed of pall rings made of metal as droplet traps.

[0321] Below the lowest dual flow tray, 16000 kg / h of compressed residue gas mixture R (2.5 bar, 150° C.) is led to a first stripping column, which rises in countercurrent to the descending extract in the stripping column. At the top of the first stripping column, a first content gas (comprising, inter alia, water and acrylic acid) is led off (16851 kg / h) and fed to a second stripping column. The composition of the first content gas is as follows:

[0322] [Table 14]

[0323] Bottom liquid is continuously withdrawn from the bottom of the first stripping column. A portion of the bottom liquid withdrawn from the first stripping column is cooled in a two-stage indirect heat exchange (the first stage is a plate heat exchanger with thermal integration against the extraction) and recycled to the top of the extraction column as the extractant. Another portion of the bottom liquid withdrawn from the first stripping column is heated to about 160° C. in an external forced circulation tube bundle flash evaporator and recycled to the bottom of the first stripping column.

[0324] The overall propene conversion (based on multiple passes) in this Example 1 is 99.86 mole % (calculated from the difference between the propene content of the chemical-grade propene fed to the production process via the initial reaction gas mixture 1 and the propene content of the off-gas stream O discharged from the production process).

[0325] The total yield of acrylic acid produced in Example 1 is 87.2 mol % (calculated from the acrylic acid as the content of glacial acrylic acid continuously fed to the storage tank and the propene as the content of chemical grade propene fed to the production process via the initial reaction gas mixture 1).

[0326] From the acrylic acid contained in the product gas mixture 2, 99.40 mol % passes into the liquid phase in the separation zone.

[0327] Production of synthesis gas from the off-gas of this Example 1 In this example 1, the off-gas released from the process carried out for the production of the target product acrylic acid can be used directly as a raw material for the production of synthesis gas, which is advantageously carried out by applying the following reactions a) to c) in combination, which reactions are usually carried out at high temperature and pressure: a) heterogeneously catalyzed steam reforming of the hydrocarbons contained in the off-gas; b) heterogeneously catalyzed hydrogenation of the CO2 contained in the off-gas (e.g., via the hydrogen formed in reaction a); and c) Oxidation of organic by-products and target product components still contained in the off-gas via the residual molecular oxygen and / or water contained in the off-gas.

[0328] Preferably, a catalyst is used that is capable of catalyzing more than one of the above reactions a), b) and c), best all three. Such catalysts are disclosed, for example, in WO 2015 / 135968 (e.g., Example 1 of said WO).

[0329] On a technical scale, such production is usually carried out in externally heated tubular reformers (see, for example, Ullmann's Encyclopedia of Industrial Chemistry, Fifth Completely Revised Edition, Volume A12 (Formamides to Hexamethylendiamine), VCH Verlagsgesellschaft mbH, D-6940 Weinheim, pages 192-194, section 2.2.3.).

[0330] On a laboratory scale, a tubular reformer can be represented by a single externally heated reaction tube that can be designed, loaded with catalyst, and operated as follows.

[0331] A technically advantageous tubular laboratory steam reformer is an externally heated reactor tube made of INCOLOY® Alloy 800 H (VDM material company no. 1.4876). Its total length is 100 mm, its internal diameter is 21 mm and its wall thickness is 2.87 mm. At the bottom of the tubular reactor, a sieve ((square) mesh size no. 70 (nominal size opening = 212 μm according to US ASTM standard sieve specification E-11)) is placed on a holder (sieve and holder are also prepared from INCOLOY® alloy 800H). At the top of the sieve (located at a height of 10 mm from the bottom) a first fixed bed (length = 10 mm) of inert steatite (C220 steatite, CeramTec GmbH, D-73207 Plochingen) is placed having particles in the size range of 300 μm to 500 μm (measured via sieving; (square) mesh size no. 40 sieve (nominal size opening = 425 μm, according to US ASTM standard sieve specification E-11)). Before filling the still empty tube volume with catalytically active particles coming from the bottom of the tubular reactor, a single sheathed (sheath material: Inconel® alloy 600 (VDM material company no. 2.4816)) Pt-Pt / Rh thermocouple (diameter=3 mm) is introduced (in the center of the tube) in such a way that its single metering point is located 30 mm above the top of the first fixed bed of inert steatite particles (to determine the actual temperature value compared to a target value in order to control the temperature in the catalytic bed). Subsequently, a fixed bed (length = 60 mm) of reformer catalyst particles (particle composition and preparation as for catalyst no. 1 in table 1 of WO 2015 / 135968; particle size in the range of 300 μm to 500 μm (measured through sieving; (square) mesh size no. 40 sieve (nominal size opening = 425 μm, according to US ASTM standard sieve specification E-11)) is placed on top of the first fixed bed of inert steatite particles. A second fixed bed (length = 20 mm) of the same inert steatite particles is placed on top of the fixed bed of catalytically active reformer catalyst particles. The tubular reactor is heated from its outside by an electric jacket heater over its entire length of 100 mm. Each gas stream flows from top to bottom through the reactor tube.The second fixed bed of inert steatite particles serves the purpose of bringing the temperature of the flowing gas stream to the desired target temperature (tested by thermocouple) by entering the catalytically active fixed bed section.

[0332] Before starting the synthesis gas production, the reformer catalyst particles are activated with a stream of diluted hydrogen (5% by volume H2 and 95% by volume N2; both exhibiting a purity of 4.0, respectively (purity ≥ 99.99% by volume; ≤ 50 ppm by volume O2 and ≤ 30 ppm by volume H2O) at a temperature of 450 °C (always the temperature at the metering point of the thermocouple) for a period of 10 hours. The hourly space velocity of molecular hydrogen over the fixed catalyst bed (only the catalytically active section in the reactor tube, not including the two inactive sections) is 190 (l(STP) / l·h). This corresponds to a total gas flow of 1.33 l(STP) / min. The pressure at the inlet to the tubular reactor is 20 bar. The heating of the fixed catalyst bed to 450 °C (through which the specified diluted molecular hydrogen as described above flows continuously during the heating) is carried out via application of the following temperature-programmed procedure:

[0333] Starting from ambient temperature (approximately 25° C.), the fixed bed is initially heated to 250° C. (always the temperature at the thermocouple metering point) at a rate of 10° C. per minute. The temperature of the fixed bed is then held at this intermediate value for 25 minutes, before it is initially increased to 350° C. at a rate of 1° C. per minute, and after maintaining this temperature for 25 minutes, it is increased to 450° C. at the same heating rate. This temperature is then maintained for 10 hours for activation purposes as already described. Once activation is complete, the fixed bed temperature is increased to the target value 950° C. suitable for the intended synthesis gas production at a heating rate of 10° C. per minute (as always the temperature at the thermocouple metering point) and maintained at this temperature for 1 hour.

[0334] Before introducing the reaction gas mixture into the steam reformer tubular reactor packed with a fixed bed, the flow of diluted hydrogen is switched to a flow of diluted steam (20% by volume H2O and 80% by volume N2 (purity 4.0)). A total gas flow of 1.33 l (STP) / min and a pressure of 20 bar and 950° C. are maintained at the tubular reactor inlet. This measure helps to avoid the formation of coke when the reaction gas enters the reaction tubes.

[0335] To initiate synthesis gas formation, 80% by volume of the N2 content of the diluted steam stream (assuming other operating conditions remain constant) is replaced by the same content (80% by volume) of off-gas.

[0336] In this Example 1, in the case of the off-gas discharged from the process carried out for the production of the target product acrylic acid, the synthesis gas leaving the bottom of the reactor tube has the following composition (the data presented are derived from the results obtained in an equivalent single tube experiment, applying the corresponding catalyst charge and similar / equivalent initial reaction gas mixture and reaction conditions):

[0337] [Table 15]

[0338] The synthesis gas produced is rich in molecular hydrogen and carbon monoxide and can therefore be used for the synthesis of chemical compounds containing C and H, and optionally O. Before such use of the still contained HO, it can be easily separated by condensation means, if so desired. Alternatively, the synthesis gas produced can be mixed with any other synthesis gas produced in any plant dedicated to the production of synthesis gas (for example, a coal gasification plant), and only the resulting mixture of synthesis gas is applied to the synthesis of the above-mentioned chemical compounds. Instead of such use of the off-gas discharged from any of the inventive methods for the production of synthesis gas, such off-gas can be mixed with other feedstocks (for example, a mixture of steam and methane) that are usually supplied in any plant dedicated solely to the production of synthesis gas (for example, a tubular methane reformer), and the resulting mixture is then supplied to such a dedicated synthesis gas production plant.

[0339] The synthesis gas produced in this Example 1 is particularly valuable so far because it contains carbon dioxide as well as molecular hydrogen and carbon monoxide. This is because, for example, methanol can be made from carbon monoxide and carbon dioxide (2H2+CO⇔CH3OH and 3H2+CO2⇔CH3OH and H2O). The preferred reaction pathway goes through carbon monoxide, with some carbon dioxide needed for kinetic reasons. The concentration ratio R C =(c H2 -c CO2 ) / c CO +c CO2 ) should generally be between 2.0 and 2.2 (see, for example, Ullmann's Encyclopedia of Industrial Chemistry, Fifth Completely Revised Edition, Volume A12 (Formamides to Hexamethylendiamine), VCH Verlagsgesellschaft mbH, D-6940 Weinheim, p. 174). The synthesis gas produced in this Example 1 has a value R of only +0.37. CHowever, this value can be easily increased to a value between 2.0 and 2.2 by adding a corresponding amount of molecular hydrogen to the synthesis gas before carrying out the relevant synthesis reaction.

[0340] Through the addition of molecular hydrogen to the off-gas discharged from the process carried out in this Example 1 for the production of the target product acrylic acid (this and the previously mentioned molecular hydrogen can advantageously bring about the propane dehydrogenation used for the production of the propene feedstock used to prepare the starting reaction gas mixture 1), before using said off-gas for the production of the synthesis gas outlined above, it is even possible to produce a synthesis gas with a negligible carbon dioxide content and containing carbon monoxide and molecular hydrogen in a molar ratio of 1:1.

[0341] Such synthesis gas is preferably applicable to the hydroformylation of olefins (especially propene), as described in WO 2002 / 000587, WO 2017 / 108878, WO 2018 / 210720 and WO 2005 / 095315. The aldehydes (especially butyraldehyde) resulting in such hydroformylation reactions can subsequently be hydrogenated to produce the corresponding alcohols (especially n-butanol) and finally esterified via reaction with acrylic acid (preferably produced by the process claimed in this application) to produce the corresponding acrylates (especially butyl acrylate). If the hydrocarbon content of the off-gas is not sufficient for the proper application of reaction a) (heterogeneously catalyzed steam reforming of the hydrocarbons contained in the off-gas), such hydrocarbons can be easily added to the off-gas before it is used for the production of synthesis gas (preferably such hydrocarbons are already part of the initial reaction gas mixture 1).

[0342] Regardless of how the off-gas produced in this Example 1 is specifically used for the production of synthesis gas, and regardless of how such synthesis gas is specifically used for the production of chemical compounds, what is common to all the procedures described above is that any "C" contained in any compound of the off-gas is no longer combusted to produce CO2, but is advantageously used for the synthesis of "C"-containing molecules.

[0343] However, the economic implementations described above assume that the relevant off-gases have only a limited content of molecular nitrogen, since such molecular nitrogen is ultimately contained in the synthesis gas produced. However, chemical synthesis processes based on synthesis gas (e.g., hydroformylation of propene) are usually not continuous processes resulting in 100% conversion of the starting materials involved (e.g., propene). The resulting product mixture therefore still contains unreacted raw materials and inert molecular nitrogen. It is naturally economically desirable to recycle said unreacted raw materials to the respective synthesis process.

[0344] However, for a continuous chemical synthesis gas process (e.g. hydroformylation) to operate successfully on an industrial scale, it is important to establish a steady state between the feed material, including any recycle, and the degree of reaction. It is therefore important to prevent excessive N2 build-up in the reaction system by recycling unreacted components. However, molecular nitrogen is difficult to separate from other compounds contained in, for example, CO, CO2, propene, H2, and others. Thus, to prevent significant N2 build-up from recycle, it is necessary to vent at least some molecular nitrogen. All this removal is more significant the higher the N2 content of the off-gas used for synthesis gas production. Any such N2 removal necessarily involves the removal of some remaining feedstock, which automatically causes some undesirable inefficiencies and economic liabilities in the process. EXAMPLES

[0345] Describes steady state (small amounts of O2 dissolved in the liquid material stream are not included (not mentioned) below)

[0346] Starting reactant gas mixture 1 is a mixture of the following feed gas streams: - 16534 kg / h of molecular oxygen (purity: 99.9999% by mass O2); - 16000 kg / h of polymer grade propene (99.48% by mass of propene and 0.52% by mass of propane); - a recycled gas stream P of 203,356 kg / h; and - 158.9 kg / h of CH4 (purity: 99.999% by mass).

[0347] The temperature of the initial reaction gas mixture 1 is 200° C. and its pressure is 2.434 bar. The composition of the initial reaction gas mixture 1 fed through the first stage tube bundle reactor of reaction tubes is as follows:

[0348] [Table 16]

[0349] temperature T A is 329°C. Temperature T B is 336°C.

[0350] The temperature of the catalytically active part of the fixed catalyst bed in the reaction tube is in the range of 326°C to 403°C.

[0351] The hourly space velocity of propene over the fixed catalyst bed (section 2 + section 3 in the reactor tube) is 172 (l(STP) / l·h).

[0352] The single-pass propene conversion was 95.8 mol % with a concomitant selectivity for the formation of acrolein S A is 91.7 mol%.

[0353] The flow of 236048 kg / h of product gas mixture 1 leaving the first reaction stage has the following composition at a temperature of 360° C. and a pressure of 1.834 bar:

[0354] [Table 17]

[0355] The temperature of the product gas mixture leaving the first reaction stage (product gas mixture 1) is reduced by indirect cooling in an aftercooler. The thus cooled product gas mixture 1 stream (temperature=250° C.) is mixed with a stream of 3792.4 kg / h of secondary molecular oxygen (purity: 99.9999% by weight O2; temperature: 165.00° C.; pressure: 2.50 bar).

[0356] The composition of the thus produced starting reaction gas mixture 2 (temperature: 249.0° C.; pressure: 1.809 bar), fed through the reaction tubes in the second-stage tube bundle reactor, of 239841 kg / h is as follows:

[0357] [Table 18]

[0358] temperature T C is 268°C. Temperature T D is 273°C.

[0359] The temperature of the catalytically active portion of the fixed catalyst bed in the reaction tube is in the range of 266°C to 305°C.

[0360] The hourly space velocity of acrolein over the fixed catalyst bed (sections 2 and 3 in the reactor tube) is 138 (l(STP) / l·h).

[0361] The single pass acrolein conversion was 99.5 mol % and the associated selectivity for the formation of acrylic acid, S, based on the converted propene, was evaluated over both reaction stages. AA is 89.8 mol%.

[0362] Leaving the second stage reactor (product gas mixture 2) and at its inlet into the separation zone, the product gas mixture (239841 kg / h, T=250° C., P=1.509 bar) is cooled to a temperature of 120° C. in a spray cooler (quench 1).

[0363] The composition of product gas mixture 2 is as follows:

[0364] [Table 19]

[0365] The liquid to be used for direct cooling of the product gas mixture 2 (quench liquid 1) is a part (516483 kg / h, T=116.2° C.) of the bottom liquid withdrawn from the bottom of the condensation column described below in a total amount of 519893 kg / h (T=116.2° C.). The content of this bottom liquid is:

[0366] [Table 20]

[0367] The mixture of product gas mixture cooled to 120° C. and unevaporated quench liquid 1, which has the same temperature as a result of direct cooling, is thus led to the bottom of the condensation column. The pressure in the bottom space and in the quench 1 is 1.479 bar. The other part of the bottom liquid (3410 kg / h) withdrawn from the bottom of the condensation column is made up to the second stripping column and fed to its central section.

[0368] The energy required for cracking the diacrylic acid and polyacrylic acid (both Michael adducts) contained in the bottom liquid to acrylic acid is replenished to the second stripping column by means of an external forced circulation three-flow tube bundle flash evaporator, which is fed with the bottom liquid of the second stripping column. As it flows through the heat exchanger tubes, the bottom liquid is heated and the main amount of it is then recycled to the bottom of the second stripping column. A small amount of the total amount of bottom liquid led through the heat exchanger is branched, degassed, diluted with methanol and sent to residue incineration.

[0369] In addition, the first content gas withdrawn from the top of the first stripping column is fed to the bottom of the second stripping column. The second content gas is withdrawn from the top of the second stripping column in an amount of 18868 kg / h (temperature=65° C., pressure=1.40 bar) and fed to the bottom of the quench 1 and / or condensation column. The content of the second content gas is as follows:

[0370] [Table 21]

[0371] 88253 kg / h of the high boiler fraction is led from the first collecting tray of the condensation column to the bottom space arranged below the first collecting tray. The high boiler fraction has the following content at a temperature of 103.2° C. and a pressure of approximately 1.479 bar:

[0372] [Table 22]

[0373] The bottom temperature is 116.2° C. and the bottom pressure (at liquid level) is 1.479 bar.

[0374] From the second collecting tray, 254449 kg / h of crude acrylic acid is continuously withdrawn as a first side draw at a temperature of 100.8° C. and 1.443 bar, which has the following content:

[0375] [Table 23]

[0376] The 43,857 kg / h of crude acrylic acid withdrawn from the second collecting tray is obtained by further purification by crystallization of the withdrawn acrylic acid, and is recycled to the condensation column via the respective inserted tubes immediately below the second collecting tray, together with the mother liquor (71,694 kg / h) which is heated to 90°C by indirect heat exchange with the withdrawn crude acrylic acid and steam as a heat medium, to the dual flow tray following below the second collecting tray.

[0377] 115000 kg / hr of crude acrylic acid withdrawn from the second collection tray is recycled directly onto the second collection tray via spray nozzles at a temperature of 100.8° C. to keep this section wet and prevent contamination.

[0378] 95592 kg / h of crude acrylic acid withdrawn from the second collecting tray is cooled by multi-stage indirect heat exchange to a temperature of 29° C. About 1000 kg / h of deionized water is then added to the cooled acrylic acid.

[0379] The resulting mixture is cooled and subsequently crystallized in a cooled disk crystallizer, and the resulting crystal suspension is purified in a hydraulic melt wash column as described in WO 2008 / 090190, pages 49 and 50. From the melt circuit, stabilized by the addition of a solution of MEHQ in glacial acrylic acid, 23898 kg / h of the desired separated acrylic acid are withdrawn as glacial acrylic acid, which has the following contents:

[0380] [Table 24]

[0381] Inhibitor solution 1 is prepared at 25° C. by dissolving 11 kg / h of PTZ (phenothiazine) in 904 kg / h of the abovementioned glacial acrylic acid. The remaining stream of glacial acrylic acid, continuously withdrawn from the melting circuit, is continuously fed to a storage tank.

[0382] The mother liquor removed in the wash column is first conducted to a heatable collecting vessel and from there to a tank from which it is heated to 90° C. by thermal integration (as already described) and recycled to the condensation column with a mass flow rate of 71694 kg / h (as a mixed stream (94.1° C., 1.10 bar)) together with 43857 kg / h of crude acrylic acid withdrawn on the second collecting tray of the condensation column in the upper part of the series of dual flow trays below the second collecting tray.

[0383] The composition of the recycled mother liquor is as follows:

[0384] [Table 25]

[0385] 841201 kg / hr of acid water is withdrawn as a second side draw from the third collection tray at a temperature of 58.5° C. and a pressure of 1.229 bar.

[0386] Acid water has the following content:

[0387] [Table 26]

[0388] The withdrawn 31515 kg / h of acid water (58.5° C.) is recycled to the top Soman tray together with 35 kg / h of inhibitor solution 1 (25° C.) and 28 kg / h of molten MEHQ (T=80° C.). 880 kg / h of inhibitor solution 1 is recycled to the single-flow Soman tray arrangement (temperature 25° C.) after approximately two-thirds of the length of this Soman separation section (counting from its bottom).

[0389] Pulled out 458.73m 3 / hour of acid water, at a temperature of 28.2°C, is recycled to the centre of an equidistant arrangement of valve trays, which are described in more detail below (cooling is effected by means of multi-stage indirect heat exchange).

[0390] The withdrawn 325000 kg / h of acid water are recycled at a temperature of 23.3° C. to the top of the above-mentioned equidistant arrangement of valve trays (cooling is carried out together with the above-mentioned amount of acid water by means of multi-stage indirect heat exchange; the final cooling stage from 28.2° C. to 23.3° C. is carried out thermally and with the aid of heat integration (liquid chemical-grade propene is used as coolant and is evaporated simultaneously; the gaseous propene produced is subsequently used to constitute the starting reaction gas mixture 1 for the first stage of the heterogeneously catalyzed gas-phase partial oxidation of propene).

[0391] The withdrawn 9686 kg / hour of acid water is fed to an extraction column for the purpose of subsequent extraction.

[0392] Above the third collecting tray in the condensation column, a series of two-flow valve trays are fixed in an equidistant arrangement. The pressure at the top of the condensation column is 1.17 bar. At the top of the condensation column, 222658 kg / h of residual gas mixture R leaves the separation column at a temperature of 24° C. and with the following content:

[0393] [Table 27]

[0394] In the indirect heat exchanger, the gas mixture R is heated to 32° C., of which 3303 kg / h are discharged as off-gas stream O from the production process.

[0395] 219356 kg / h of said gas mixture R are compressed by means of a compressor to a pressure of 2.5 bar, raising its temperature to approximately 150° C. 16000 kg / h of it are fed to a first stripping column in the separation zone for stripping the extract from the acid water extraction. The remaining 203356 kg / h of the compressed gas mixture (as cycle gas) are recycled as gas stream P to the first reaction stage 1 as feed stream for preparing the stream of starting reaction gas mixture 1. The sum of the mass flows of the off-gas stream O and the cycle gas stream P constitutes the total mass flow M of the remaining gas mixture R, which is continuously led off from the separation zone. It contains, as pure compounds, 99.99 mol % of the total molar amount of those compounds contained in the product gas mixture 2, which is continuously transferred to the separation zone, having a boiling point below 250 K at a working pressure of 1 bar.

[0396] The mass flow rate of the recycle gas stream P is 98.40% of the total mass flow rate M of the remaining gas mixture R, which is continuously led off from the separation zone.

[0397] 9686 kg / h of acid water to be extracted (temperature=58.5° C.) are fed to the extraction column below the bottom packing via a tubular distributor with suitable passage openings. Above the top packing of the extraction column, a mixture of fresh Palatinol® A and extraction solvent recycled from the first stripping column and previously stripped free therein (ratio of the respective mass flows 1:240) (temperature=50° C.) is introduced at a mass flow rate essentially equal to the 9686 kg / h of acid water. The recycled extraction solvent has the following content:

[0398] [Table 28]

[0399] The specific mass of acid water (at temperature = 58.5°C) is 1013.4 kg / m 3 The extraction solvent is likewise introduced through a tubular distributor with suitable passage openings. The acid water forms a continuous phase, while the extraction solvent forms a dispersed phase in the form of droplets (liquid particle diameter is in the range of 2-5 mm), which descends within the aqueous phase. At the top of the extraction column a raffinate is withdrawn. This raffinate is sent to incineration. The extract is withdrawn from the bottom of the extraction column. It contains the acrylic acid extracted from the acid water. The entire extract is led to the top of the first stripping column. The extract is preheated to 95 ° C by indirect heat exchange in a plate heat exchanger. The heat carrier is the bottom liquid withdrawn from the first stripping column.

[0400] Below the lowest dual flow tray, 16000 kg / h of compressed residue gas mixture R (2.5 bar, 150° C.) is led to a first stripping column, which rises in countercurrent to the descending extract in the stripping column. At the top of the first stripping column, a first contained gas (containing, inter alia, water and acrylic acid) is led off (16868 kg / h) and fed to a second stripping column. The composition of the first contained gas is as follows:

[0401] [Table 29]

[0402] Bottom liquid is continuously withdrawn from the bottom of the first stripping column. A portion of the bottom liquid withdrawn from the first stripping column is cooled in a two-stage indirect heat exchange (the first stage is a plate heat exchanger with thermal integration against the extraction) and recycled to the top of the extraction column as the extractant. Another portion of the bottom liquid withdrawn from the first stripping column is heated to about 160° C. in an external forced circulation tube bundle flash evaporator and recycled to the bottom of the first stripping column.

[0403] The overall propene conversion (based on multiple passes) in this Example 1 is 99.93 mole % (calculated from the difference between the propene content of the chemical-grade propene fed to the production process via the initial reaction gas mixture 1 and the propene content of the off-gas stream O discharged from the production process).

[0404] The total yield of acrylic acid produced in Example 1 is 87.4 mol % (calculated from the acrylic acid as the glacial acrylic acid content continuously fed to the storage tank and the propene as the chemical grade propene content fed to the production process via the initial reaction gas mixture 1).

[0405] From the acrylic acid contained in the product gas mixture 2, 99.36 mol % passes into the liquid phase in the separation zone.

[0406] Production of synthesis gas from off-gas in this Example 2 In this Example 2, the off-gas discharged from the process carried out for the production of the target product acrylic acid can be used as is as a raw material for the production of synthesis gas. On a laboratory scale, this production is advantageously carried out in the same manner as detailed in Example 1 herein. The only difference is that, as a raw material for the production process, the off-gas produced in Example 1 herein is replaced by the off-gas produced in this Example 2 (all other operating conditions are kept as in Example 1). The synthesis gas flowing out from the bottom of the reactor tube has the following composition:

[0407] [Table 30]

[0408] The R of this synthetic gas C The value is -0.27.

[0409] All other statements made in Example 1 herein apply accordingly. EXAMPLES

[0410] Describes steady state (small amounts of O2 dissolved in the liquid material stream are not included (not mentioned) below)

[0411] Starting reactant gas mixture 1 is a mixture of the following feed gas streams: - 17494 kg / h of molecular oxygen (purity: 99.9999% by mass O2); - 16000 kg / h of chemical grade propene (94.77% by mass propene and 5.23% by mass propane); - 63462 kg / h of recycled gas stream P; and - 27228 kg / h of CH4 (purity: 99.999% by mass).

[0412] The temperature of the initial reaction gas mixture 1 is 200° C. and its pressure is 2.434 bar. The composition of the initial reaction gas mixture 1 fed through the first stage tube bundle reactor of reaction tubes is as follows:

[0413] [Table 31]

[0414] temperature T A is 326°C. Temperature T B is 335°C.

[0415] The temperature of the catalytically active portion of the fixed catalyst bed in the reaction tube is in the range of 325°C to 400°C.

[0416] The hourly space velocity of propene over the fixed catalyst bed (section 2 + section 3 in the reactor tube) is 162 (l(STP) / l·h).

[0417] The single-pass propene conversion was 95.8 mol % with a concomitant selectivity for the formation of acrolein S A is 91.7 mol%.

[0418] The flow of 124184 kg / h of product gas mixture 1 leaving the first reaction stage has the following composition at a temperature of 360° C. and a pressure of 1.834 bar:

[0419] [Table 32]

[0420] The temperature of the product gas mixture leaving the first reaction stage (product gas mixture 1) is reduced by indirect cooling in an aftercooler. The thus cooled stream of product gas mixture 1 (temperature = 250 °C) is mixed with a stream of 3554 kg / h of secondary molecular oxygen (purity: 99.9999% by mass O2; temperature: 165.00 °C; pressure: 2.50 bar).

[0421] The composition of the thus produced starting reaction gas mixture 2 (temperature: 249.1° C.; pressure: 1.809 bar) of 127737 kg / h, fed through the reaction tubes in the second-stage tube bundle reactor, is as follows:

[0422] [Table 33]

[0423] temperature T C is 266°C. Temperature T D is 271°C.

[0424] The temperature of the catalytically active part of the fixed catalyst bed in the reaction tube is in the range of 265°C to 301°C.

[0425] The hourly space velocity of acrolein over the fixed catalyst bed (sections 2 and 3 in the reactor tube) is 130 (l(STP) / l·h).

[0426] The single pass acrolein conversion was 99.5 mol % and the associated selectivity for the formation of acrylic acid, S, based on the converted propene, was evaluated over both reaction stages. AA is 89.7 mol%.

[0427] Leaving the second stage reactor (product gas mixture 2) and at its inlet into the separation zone, the product gas mixture (127737 kg / h, T=250° C., P=1.509 bar) is cooled to a temperature of 120° C. in a spray cooler (quench 1).

[0428] The composition of product gas mixture 2 is as follows:

[0429] [Table 34]

[0430] The liquid to be used for direct cooling of the product gas mixture 2 (quench liquid 1) is a part (513405 kg / h, T=116.1° C.) of the bottom liquid withdrawn from the bottom of the condensation column described below in a total amount of 516815 kg / h (T=116.1° C.). The content of this bottom liquid is:

[0431] [Table 35]

[0432] The mixture of product gas mixture cooled to 120° C. and unevaporated quench liquid 1, which has the same temperature as a result of direct cooling, is thus led to the bottom of the condensation column. The pressure in the bottom space and in the quench 1 is 1.479 bar. The other part of the bottom liquid withdrawn from the bottom of the condensation column (3410 kg / h) is made up to the second stripping column and fed to its central section.

[0433] The energy required for cracking the diacrylic acid and polyacrylic acid (both Michael adducts) contained in the bottom liquid to acrylic acid is replenished to the second stripping column by means of an external forced circulation three-flow tube bundle type flash evaporator, which is fed to the bottom liquid of the second stripping column. As it flows through the heat exchanger tubes, the bottom liquid is heated, the main amount of which is then recycled to the bottom of the second stripping column. A small part of the total amount of bottom liquid led through the heat exchanger is branched, degassed, diluted with methanol and sent to residue incineration.

[0434] In addition, the first content gas withdrawn from the top of the first stripping column is fed to the bottom of the second stripping column. The second content gas withdrawn from the top of the second stripping column in an amount of 18860 kg / h (temperature=65° C., pressure=1.40 bar) is fed to the bottom of the quench 1 and / or condensation column. The content of the second content gas is as follows:

[0435] [Table 36]

[0436] 85403 kg / h of the high boiler fraction is led from the first collecting tray of the condensation column to the bottom space arranged below the first collecting tray. The high boiler fraction has the following content at a temperature of 102.3° C. and a pressure of approximately 1.479 bar:

[0437] [Table 37]

[0438] The bottom temperature is 116.1° C. and the bottom pressure (liquid level) is 1.479 bar.

[0439] From the second collecting tray, 252432 kg / h of crude acrylic acid is continuously withdrawn as a first side draw at a temperature of 100.0° C. and 1.443 bar, which has the following content:

[0440] [Table 38]

[0441] The 47,690 kg / h of crude acrylic acid withdrawn from the second collecting tray is obtained by further purification by crystallization of the withdrawn acrylic acid, and is recycled to the condensation column via the respective inserted tubes immediately below the second collecting tray, together with the mother liquor (67,306 kg / h) heated to 90°C by indirect heat exchange with the withdrawn crude acrylic acid and steam as a heat medium, to the dual flow tray following below the second collecting tray.

[0442] The 115,000 kg / hr crude acrylic acid withdrawn from the second collection tray is recycled directly onto the second collection tray via spray nozzles at a temperature of 100.0° C. to keep this section wet and prevent contamination.

[0443] 89741 kg / h of crude acrylic acid withdrawn from the second collecting tray is cooled by multi-stage indirect heat exchange to a temperature of 29° C. About 1000 kg / h of deionized water is then added to the cooled acrylic acid.

[0444] The resulting mixture is cooled and subsequently crystallized in a cooled disk crystallizer, and the resulting crystal suspension is purified in a hydraulic melt wash column as described in WO 2008 / 090190, pages 49 and 50. From the melt circuit, stabilized by the addition of a solution of MEHQ in glacial acrylic acid, 22435 kg / h of the desired separated acrylic acid are withdrawn as glacial acrylic acid, which has the following contents:

[0445] [Table 39]

[0446] Inhibitor solution 1 is prepared by dissolving 11 kg / h of PTZ (phenothiazine) in 904 kg / h of the abovementioned glacial acrylic acid at 25° C. The remaining stream of glacial acrylic acid, which is continuously withdrawn from the melting circuit, is continuously fed to a storage tank.

[0447] The mother liquor removed in the wash column is first conducted to a heatable collecting vessel and from there to a tank from which it is heated to 90° C. by thermal integration (as already described) and recycled to the condensation column with a mass flow rate of 67306 kg / h together with 47690 kg / h of crude acrylic acid withdrawn on the second collecting tray of the condensation column (as a mixed stream (94.2° C., 1.10 bar)) in the upper part of the series of dual flow trays below the second collecting tray.

[0448] The composition of the recycled mother liquor is as follows:

[0449] [Table 40]

[0450] 839235 kg / hr of acid water is withdrawn as a second side draw from the third collection tray at a temperature of 57.2° C. and a pressure of 1.229 bar.

[0451] Acid water has the following content:

[0452] [Table 41]

[0453] The withdrawn 31052 kg / h of acid water (57.2° C.) is recycled to the top Soman tray together with 35 kg / h of inhibitor solution 1 (25° C.) and 28 kg / h of molten MEHQ (T=80° C.). 880 kg / h of inhibitor solution 1 is recycled to the single-flow Soman tray arrangement (temperature 25° C.) after approximately two-thirds of the length of this Soman separation section (counting from its bottom).

[0454] Pulled out 458.54m 3 / hour of acid water, at a temperature of 28.2°C, is recycled to the centre of an equidistant arrangement of valve trays, which are described in more detail below (cooling is effected by means of multi-stage indirect heat exchange).

[0455] The withdrawn 325000 kg / h of acid water are recycled at a temperature of 23.3° C. to the top of the above-mentioned equidistant arrangement of valve trays (cooling is carried out together with the above-mentioned amount of acid water by means of multi-stage indirect heat exchange; the final cooling stage from 28.2° C. to 23.3° C. is carried out thermally and with the aid of heat integration (liquid chemical-grade propene is used as a coolant and is evaporated simultaneously; the gaseous propene produced is subsequently used to constitute the starting reaction gas mixture 1 for the first stage of the heterogeneously catalyzed gas-phase partial oxidation of propene).

[0456] The withdrawn 8183 kg / h of acid water is fed to an extraction column for the purpose of subsequent extraction.

[0457] Above the third collecting tray in the condensation column, a series of two-flow valve trays are fixed in an equidistant arrangement. The pressure at the top of the condensation column is 1.17 bar. At the top of the condensation column, 113512 kg / h of residual gas mixture R leaves the separation column at a temperature of 24° C. and with the following content:

[0458] [Table 42]

[0459] In the indirect heat exchanger, the gas mixture R is heated to 32° C. 34051 kg / h thereof are discharged as off-gas stream O from the production process.

[0460] 79462 kg / h of said gas mixture R are compressed by means of a compressor to a pressure of 2.5 bar, raising its temperature to approximately 150° C. 16000 kg / h of it are fed to a first stripping column in the separation zone for stripping the extract from the acid water extraction. The remaining 63462 kg / h of the compressed gas mixture (as cycle gas) are recycled as gas stream P to the first reaction stage 1 as feed stream for preparing the stream of starting reaction gas mixture 1. The sum of the mass flows of the off-gas stream O and the cycle gas stream P constitutes the total mass flow M of the remaining gas mixture R, which is continuously led off from the separation zone. It contains 99.99 mol % of the total molar amount of those compounds, which, as pure compounds, have a boiling point below 250 K at a working pressure of 1 bar and are contained in the product gas mixture 2, which is continuously transferred to the separation zone.

[0461] The mass flow rate of the recycle gas stream P is 65.08% of the total mass flow rate M of the residual gas mixture R which is continuously led off from the separation zone.

[0462] 8183 kg / h of acid water to be extracted (temperature=57.2° C.) are fed to the extraction column below the bottom packing via a tubular distributor with suitable passage openings. Above the top packing of the extraction column, a mixture of fresh Palatinol® A and extraction solvent recycled from the first stripping column and previously stripped free therein (ratio of the respective mass flows 1:240) (temperature=50° C.) is introduced at a mass flow rate essentially equal to the 8183 kg / h of acid water. The recycled extraction solvent has the following content:

[0463] [Table 43]

[0464] The specific mass of acidic water (temperature = 57.2°C) is 1014.8 kg / m 3 The extraction solvent is likewise introduced through a tubular distributor with suitable passage openings. The acid water forms a continuous phase, while the extraction solvent forms a dispersed phase in the form of droplets (liquid particle diameter is in the range of 2-5 mm), which descends within the aqueous phase. At the top of the extraction column a raffinate is withdrawn. This raffinate is sent to incineration. The extract is withdrawn from the bottom of the extraction column. It contains the acrylic acid extracted from the acid water. The entire extract is led to the top of the first stripping column. The extract is preheated to 95 ° C by indirect heat exchange in a plate heat exchanger. The heat carrier is the bottom liquid withdrawn from the first stripping column.

[0465] Below the lowest dual flow tray, 16000 kg / h of compressed residue gas mixture R (2.5 bar, 150° C.) is led to a first stripping column, which rises in countercurrent to the descending extract in the stripping column. At the top of the first stripping column, a first contained gas (containing, inter alia, water and acrylic acid) is led off (16860 kg / h) and fed to a second stripping column. The composition of the first contained gas is as follows:

[0466] [Table 44]

[0467] Bottom liquid is continuously withdrawn from the bottom of the first stripping column. A portion of the bottom liquid withdrawn from the first stripping column is cooled in a two-stage indirect heat exchange (the first stage is a plate heat exchanger with thermal integration against the extraction) and recycled to the top of the extraction column as the extractant. Another portion of the bottom liquid withdrawn from the first stripping column is heated to about 160° C. in an external forced circulation tube bundle flash evaporator and recycled to the bottom of the first stripping column.

[0468] The overall propene conversion (based on multiple passes) in this Example 1 is 98.50 mol % (calculated from the difference between the propene content of the chemical grade propene fed to the production process via the initial reaction gas mixture 1 and the propene content of the off-gas stream O discharged from the production process).

[0469] The total yield of acrylic acid produced in Example 1 is 86.1 mol% (calculated from the acrylic acid as the content of glacial acrylic acid continuously fed to the storage tank and the propene as the content of chemical grade propene fed to the production process via the initial reaction gas mixture 1).

[0470] From the acrylic acid contained in the product gas mixture 2, 99.37 mol % passes into the liquid phase in the separation zone.

[0471] Production of synthesis gas from off-gas in this Example 3 In this Example 3, the off-gas discharged from the process carried out for the production of the target product acrylic acid can be used as is as a raw material for the production of synthesis gas. On a laboratory scale, this production is advantageously carried out in the same manner as detailed in Example 1 herein. The only difference is that, as a raw material for the production process, the off-gas produced in Example 1 herein is replaced by the off-gas produced in this Example 3 (all other operating conditions are kept as in Example 1). The synthesis gas flowing out from the bottom of the reactor tube has the following composition:

[0472] [Table 45]

[0473] The R of this synthetic gas C The value is +2.44. By increasing the steam content in the steam and off-gas mixture used for syngas production, the residual content of methane in the resulting syngas can be reduced, with a corresponding increase in the amounts of carbon monoxide and hydrogen in the resulting syngas.

[0474] All other statements made in Example 1 herein apply accordingly. EXAMPLES

[0475] Describes steady state (small amounts of O2 dissolved in the liquid material stream are not included (mentioned below)) Starting reactant gas mixture 1 is a mixture of the following feed gas streams: - 16336 kg / h of a mixture of molecular oxygen (96.60% by mass) and molecular nitrogen (3.40% by mass); - 16000 kg / h of chemical grade propene (94.77% by weight propene and 5.23% by weight propane); and - a recycled gas flow P of 197861 kg / h.

[0476] The temperature of the initial reaction gas mixture 1 is 200° C. and its pressure is 2.434 bar. The composition of the initial reaction gas mixture 1 fed through the first stage tube bundle reactor of reaction tubes is as follows:

[0477] [Table 46]

[0478] temperature T A is 327°C. Temperature T B is 335°C.

[0479] The temperature of the catalytically active portion of the fixed catalyst bed in the reaction tube is in the range of 326°C to 395°C.

[0480] The hourly space velocity of propene over the fixed catalyst bed (section 2 + section 3 in the reactor tube) is 164 (l(STP) / l·h).

[0481] The single-pass propene conversion was 95.8 mol % with a concomitant selectivity for the formation of acrolein S A is 91.6 mol%.

[0482] The flow of 230197 kg / h of product gas mixture 1 leaving the first reaction stage has the following composition at a temperature of 360° C. and a pressure of 1.834 bar:

[0483] [Table 47]

[0484] The temperature of the product gas mixture (product gas mixture 1) leaving the first reaction stage is reduced by indirect cooling in an aftercooler. The stream of product gas mixture 1 thus cooled (temperature=250° C.) is mixed with a stream of a secondary mixture of 3743 kg / h of molecular oxygen (96.60% by mass) and molecular nitrogen (3.40% by mass) having a temperature of 165.00° C. and a pressure of 2.50 bar.

[0485] The composition of the thus produced starting reaction gas mixture 2 (temperature: 249.1° C.; pressure: 1.809 bar) fed through the reaction tubes in the second-stage tube bundle reactor at 233940 kg / h is as follows:

[0486] [Table 48]

[0487] temperature T C is 267°C. Temperature T D is 272°C.

[0488] The temperature of the catalytically active part of the fixed catalyst bed in the reaction tube is in the range of 265°C to 303°C.

[0489] The hourly space velocity of acrolein over the fixed catalyst bed (sections 2 and 3 in the reactor tube) is 132 (l(STP) / l·h).

[0490] The single pass acrolein conversion was 99.5 mol % and the associated selectivity for the formation of acrylic acid, S, based on the converted propene, was evaluated over both reaction stages.AA is 89.8 mol%.

[0491] Leaving the second stage reactor (product gas mixture 2) and at its inlet into the separation zone, the product gas mixture (233940 kg / h, T=250° C., P=1.509 bar) is cooled to a temperature of 120° C. in a spray cooler (quench 1).

[0492] The composition of product gas mixture 2 is as follows:

[0493] [Table 49]

[0494] The liquid to be used for direct cooling of the product gas mixture 2 (quench liquid 1) is a part (522005 kg / h, T=116.4° C.) of the bottom liquid withdrawn from the bottom of the condensation column described below in a total amount of 525415 kg / h (T=116.4° C.). The content of this bottom liquid is:

[0495] [Table 50]

[0496] The mixture of product gas mixture cooled to 120° C. and unevaporated quench liquid 1, which has the same temperature as a result of direct cooling, is thus led to the bottom of the condensation column. The pressure in the bottom space and in the quench 1 is 1.479 bar. The other part of the bottom liquid withdrawn from the bottom of the condensation column (3410 kg / h) is made up to the second stripping column and fed to its central section.

[0497] The energy required for cracking the diacrylic acid and polyacrylic acid (both Michael adducts) contained in the bottom liquid to acrylic acid is replenished to the second stripping column by means of an external forced circulation three-flow tube bundle flash evaporator, which is fed with the bottom liquid of the second stripping column. As it flows through the heat exchanger tubes, the bottom liquid is heated, the main amount of which is then recycled to the bottom of the second stripping column. A small portion of the total amount of bottom liquid led through the heat exchanger is branched, degassed, diluted with methanol and sent to residue incineration.

[0498] In addition, the first content gas withdrawn from the top of the first stripping column is fed to the bottom of the second stripping column. The second content gas withdrawn from the top of the second stripping column in an amount of 18858 kg / h (temperature=65° C., pressure=1.40 bar) is fed to the bottom of the quench 1 and / or condensation column. The content of the second content gas is as follows:

[0499] [Table 51]

[0500] 95593 kg / h of the high boiler fraction is led from the first collecting tray of the condensation column to the bottom space arranged below the first collecting tray. The high boiler fraction has the following content at a temperature of 104.7° C. and a pressure of approximately 1.479 bar:

[0501] [Table 52]

[0502] The bottom temperature is 116.4° C. and the bottom pressure (liquid level) is 1.479 bar.

[0503] From the second collecting tray, 259026 kg / h of crude acrylic acid is continuously withdrawn as a first side draw at a temperature of 102.4° C. and 1.443 bar, which has the following content:

[0504] [Table 53]

[0505] The 52939 kg / h of crude acrylic acid withdrawn from the second collecting tray is obtained by further purification by crystallization of the withdrawn acrylic acid, and is recycled to the condensation column via the respective inserted tubes immediately below the second collecting tray, together with the mother liquor (68315 kg / h) which is heated to 90°C by indirect heat exchange with the withdrawn crude acrylic acid and steam as a heat medium, to the dual flow tray following below the second collecting tray.

[0506] 115000 kg / hr of crude acrylic acid withdrawn from the second collection tray is recycled directly onto the second collection tray via spray nozzles at a temperature of 102.4° C. to keep this section wet and prevent contamination.

[0507] 91087 kg / h of crude acrylic acid withdrawn from the second collecting tray is cooled by multi-stage indirect heat exchange to a temperature of 29° C. About 1000 kg / h of deionized water is added to the cooled acrylic acid.

[0508] The resulting mixture is cooled and subsequently crystallized in a cooled disk crystallizer, and the resulting crystal suspension is purified in a hydraulic melt wash column as described in WO 2008 / 090190, pages 49 and 50. From the melt circuit, stabilized by the addition of a solution of MEHQ in glacial acrylic acid, 22772 kg / h of the desired separated acrylic acid are withdrawn as glacial acrylic acid, which has the following contents:

[0509] [Table 54]

[0510] Inhibitor solution 1 is prepared by dissolving 11 kg / h of PTZ (phenothiazine) in 904 kg / h of the abovementioned glacial acrylic acid at 25° C. The remaining stream of glacial acrylic acid, which is continuously withdrawn from the melting circuit, is continuously fed to a storage tank.

[0511] The mother liquor removed in the wash column is first conducted to a heatable collecting vessel and from there to a tank from which it is heated to 90° C. by thermal integration (as already described) and recycled to the condensation column with a mass flow rate of 68,315 kg / h (as a mixed stream (95.4° C., 1.10 bar)) together with 52,939 kg / h of crude acrylic acid withdrawn on the second collecting tray of the condensation column in the upper part of the series of dual flow trays below the second collecting tray.

[0512] The composition of the recycled mother liquor is as follows:

[0513] [Table 55]

[0514] 841777 kg / hr of acid water is withdrawn as a second side draw from the third collection tray at a temperature of 59.2° C. and a pressure of 1.229 bar.

[0515] Acid water has the following content:

[0516] [Table 56]

[0517] The withdrawn 32568 kg / h of acid water (59.2° C.) is recycled to the top Soman tray together with 35 kg / h of inhibitor solution 1 (25° C.) and 28 kg / h of molten MEHQ (T=80° C.). 880 kg / h of inhibitor solution 1 is recycled to the single-flow Soman tray arrangement (temperature 25° C.) after approximately two-thirds of the length of this Soman separation section (counting from its bottom).

[0518] Pulled out 458.79m 3 / hour of acid water, at a temperature of 28.2°C, is recycled to the centre of an equidistant arrangement of valve trays, which are described in more detail below (cooling is effected by means of multi-stage indirect heat exchange).

[0519] The withdrawn 325000 kg / h of acid water are recycled at a temperature of 23.3° C. to the top of the above-mentioned equidistant arrangement of valve trays (cooling is carried out together with the above-mentioned amount of acid water by means of multi-stage indirect heat exchange; the final cooling stage from 28.2° C. to 23.3° C. is carried out thermally and with the aid of heat integration (liquid chemical-grade propene is used as coolant and is evaporated simultaneously; the gaseous propene produced is subsequently used to constitute the starting reaction gas mixture 1 for the first stage of the heterogeneously catalyzed gas-phase partial oxidation of propene).

[0520] The withdrawn 9209 kg / h of acid water is fed to an extraction column for the purpose of the subsequent extraction.

[0521] Above the third collecting tray in the condensation column, a series of two-flow valve trays are fixed in an equidistant arrangement. The pressure at the top of the condensation column is 1.17 bar. At the top of the condensation column, 218351 kg / h of residual gas mixture R leaves the separation column at a temperature of 24° C. and with the following content:

[0522] [Table 57]

[0523] In an indirect heat exchanger, the gas mixture R is heated to 32° C. 4490 kg / h thereof are discharged as off-gas stream O from the production process.

[0524] 213861 kg / h of said gas mixture R are compressed by means of a compressor at a pressure of 2.5 bar, raising its temperature to approximately 150° C. 16000 kg / h of it are fed to a first stripping column in the separation zone for stripping the extract from the acid water extraction. The remaining 197861 kg / h of the compressed gas mixture (as cycle gas) are recycled as gas stream P to the first reaction stage 1 as feed stream for preparing the stream of starting reaction gas mixture 1. The sum of the mass flows of the off-gas stream O and the cycle gas stream P constitutes the total mass flow M of the remaining gas mixture R continuously led off from the separation zone. It contains 99.99 mol % of the total molar amount of those compounds contained in the product gas mixture 2 continuously transferred to the separation zone, as pure compounds, having a boiling point below 250 K at a working pressure of 1 bar.

[0525] The mass flow rate of the recycle gas stream P is 97.78% of the total mass flow rate M of the remaining gas mixture R which is continuously led off from the separation zone.

[0526] 9209 kg / h of acid water to be extracted (temperature=59.2° C.) are fed to the extraction column below the bottom packing via a tubular distributor with suitable passage openings. Above the top packing of the extraction column, a mixture of fresh Palatinol® A and extraction solvent recycled from the first stripping column and previously stripped free therein (ratio of the respective mass flows 1:240) (temperature=50° C.) is introduced at a mass flow rate essentially equal to the 9209 kg / h of acid water. The recycled extraction solvent has the following content:

[0527] [Table 58]

[0528] The specific mass of acidic water (at temperature = 59.2°C) is 1012.7 kg / m 3 The extraction solvent is likewise introduced through a tubular distributor with suitable passage openings. The acid water forms a continuous phase, while the extraction solvent forms a dispersed phase in the form of droplets (liquid particle diameter is in the range of 2-5 mm), which descends within the aqueous phase. At the top of the extraction column a raffinate is withdrawn. This raffinate is sent to incineration. The extract is withdrawn from the bottom of the extraction column. It contains the acrylic acid extracted from the acid water. The entire extract is led to the top of the first stripping column. The extract is preheated to 95 ° C by indirect heat exchange in a plate heat exchanger. The heat carrier is the bottom liquid withdrawn from the first stripping column.

[0529] Below the lowest dual flow tray, 16000 kg / h of compressed residue gas mixture R (2.5 bar, 150° C.) is led to a first stripping column, which rises in countercurrent to the descending extract in the stripping column. At the top of the first stripping column, a first contained gas (containing, inter alia, water and acrylic acid) is led off (16858 kg / h) and fed to a second stripping column. The composition of the first contained gas is as follows:

[0530] [Table 59]

[0531] Bottom liquid is continuously withdrawn from the bottom of the first stripping column. A portion of the bottom liquid withdrawn from the first stripping column is cooled in a two-stage indirect heat exchange (the first stage is a plate heat exchanger with thermal integration against the extraction) and recycled to the top of the extraction column as the extractant. Another portion of the bottom liquid withdrawn from the first stripping column is heated to about 160° C. in an external forced circulation tube bundle flash evaporator and recycled to the bottom of the first stripping column.

[0532] The overall propene conversion (based on multiple passes) in this Example 1 is 99.90 mole % (calculated from the difference between the propene content of the chemical grade propene fed to the production process via the initial reaction gas mixture 1 and the propene content of the off-gas stream O discharged from the production process).

[0533] The total yield of acrylic acid produced in Example 1 is 87.4 mol % (calculated from the acrylic acid as the glacial acrylic acid content continuously fed to the storage tank and the propene as the chemical grade propene content fed to the production process via the initial reaction gas mixture 1).

[0534] From the product gas mixture 2, which is contained in acrylic acid, 99.38 mol % passes into the liquid phase in the separation zone.

[0535] Production of synthesis gas from off-gas in this Example 4 In this Example 4, the off-gas discharged from the process carried out for the production of the target product acrylic acid can be used as is as a raw material for the production of synthesis gas. On a laboratory scale, this production is advantageously carried out in the same manner as detailed in Example 1 herein. The only difference is that, as a raw material for the production process, the off-gas produced in Example 1 herein is replaced by the off-gas produced in this Example 4 (all other operating conditions are kept as in Example 1). The synthesis gas flowing out from the bottom of the reactor tube has the following composition:

[0536] [Table 60]

[0537] The R of this synthetic gas CThe value is +0.52. By increasing the steam content in the mixture of steam and off-gas used for synthesis gas production, the residual content of propane in the synthesis gas produced can be reduced, with a corresponding increase in the amount of carbon monoxide and hydrogen in the synthesis gas produced. Due to the increase in volume during synthesis gas production (corresponding to an increase in the number of molecules), the concentration of molecular nitrogen at the point of transition from off-gas to synthesis gas drops sharply to a level that is still economically acceptable from the point of view of the subsequent use of the synthesis gas produced. All other statements made in Example 1 of this specification apply accordingly.

[0538] Comparative Example Describes steady state (small amounts of O2 dissolved in the liquid material stream are not included (not mentioned) below)

[0539] Starting reactant gas mixture 1 is a mixture of the following feed gas streams: - 87114 kg / h of a mixture of molecular oxygen (21.00% by mass) and molecular nitrogen (79.00% by mass) simulating ambient air; - 16000 kg / h of chemical grade propene (94.77% by weight propene and 5.23% by weight propane); and - A recycled gas flow P of 77110 kg / h.

[0540] The temperature of the initial reaction gas mixture 1 is 200° C. and its pressure is 2.434 bar. The composition of the initial reaction gas mixture 1 fed through the first stage tube bundle reactor of reaction tubes is as follows:

[0541] [Table 61]

[0542] temperature T A is 326°C. Temperature T B is 335°C.

[0543] The temperature of the catalytically active part of the fixed catalyst bed in the reaction tube is in the range of 325°C to 400°C.

[0544] The hourly space velocity of propene over the fixed catalyst bed (section 2 + section 3 in the reactor tube) is 161 (l(STP) / l·h).

[0545] The single-pass propene conversion was 95.5 mol % with a concomitant selectivity for the formation of acrolein S A is 92.0 mol%.

[0546] The flow of 180224 kg / h of product gas mixture 1 leaving the first reaction stage has the following composition at a temperature of 360° C. and a pressure of 1.834 bar:

[0547] [Table 62]

[0548] The temperature of the product gas mixture leaving the first reaction stage (product gas mixture 1) is reduced by indirect cooling in an aftercooler. The stream of product gas mixture 1 thus cooled (temperature=250° C.) is mixed with a stream of 18984 kg / h of a secondary mixture of molecular oxygen (21.00% by weight) and molecular nitrogen (79.00% by weight) having a temperature of 165.00° C. and a pressure of 2.50 bar.

[0549] The composition of the thus produced starting reaction gas mixture 2 (temperature: 242.9° C.; pressure: 1.809 bar) of 199208 kg / h, fed through the reaction tubes in the second-stage tube bundle reactor, is as follows:

[0550] [Table 63]

[0551] temperature T C is 265°C. Temperature T D is 271°C.

[0552] The temperature of the catalytically active part of the fixed catalyst bed in the reaction tube is in the range of 264°C to 302°C.

[0553] The hourly space velocity of acrolein over the fixed catalyst bed (sections 2 and 3 in the reactor tube) is 129 (l(STP) / l·h).

[0554] The single pass acrolein conversion was 99.5 mol % and the associated selectivity for the formation of acrylic acid, S, based on the converted propene, was evaluated over both reaction stages. AA is 89.8 mol%.

[0555] Leaving the second stage reactor (product gas mixture 2) and at its inlet into the separation zone, the product gas mixture (199208 kg / h, T=250° C., P=1.509 bar) is cooled to a temperature of 120° C. in a spray cooler (quench 1).

[0556] The composition of product gas mixture 2 is as follows:

[0557] [Table 64]

[0558] The liquid to be used for direct cooling of the product gas mixture 2 (quench liquid 1) is a part (498518 kg / h, T=116.1° C.) of the bottom liquid withdrawn from the bottom of the condensation column described below in a total amount of 501928 kg / h (T=116.1° C.). The content of this bottom liquid is:

[0559] [Table 65]

[0560] The mixture of product gas mixture cooled to 120° C. and unevaporated quench liquid 1, which has the same temperature as a result of direct cooling, is thus led to the bottom of the condensation column. The pressure in the bottom space and in the quench 1 is 1.479 bar. The other part of the bottom liquid withdrawn from the bottom of the condensation column (3410 kg / h) is made up to the second stripping column and fed to its central section.

[0561] The energy required for cracking the diacrylic acid and polyacrylic acid (both Michael adducts) contained in the bottom liquid to acrylic acid is replenished to the second stripping column by means of an external forced circulation three-flow tube bundle flash evaporator, which is fed with the bottom liquid of the second stripping column. As it flows through the heat exchanger tubes, the bottom liquid is heated, the main amount of which is then recycled to the bottom of the second stripping column. A small portion of the total amount of bottom liquid led through the heat exchanger is branched, degassed, diluted with methanol and sent to residue incineration.

[0562] In addition, the first content gas withdrawn from the top of the first stripping column is fed to the bottom of the second stripping column. The second content gas withdrawn from the top of the second stripping column in an amount of 18875 kg / h (temperature=65° C., pressure=1.40 bar) is fed to the bottom of the quench 1 and / or condensation column. The content of the second content gas is as follows:

[0563] [Table 66]

[0564] 67962 kg / h of the high boiler fraction is led from the first collecting tray of the condensation column to the bottom space arranged below the first collecting tray. The high boiler fraction has the following content at a temperature of 99.0° C. and a pressure of approximately 1.479 bar:

[0565] [Table 67]

[0566] The bottom temperature is 116.1° C. and the bottom pressure (liquid level) is 1.479 bar.

[0567] From the second collecting tray, 236076 kg / h of crude acrylic acid is continuously withdrawn as a first side draw at a temperature of 96.8° C. and 1.443 bar, which has the following content:

[0568] [Table 68]

[0569] The 32,426 kg / h of crude acrylic acid withdrawn from the second collecting tray is obtained by further purification by crystallization of the withdrawn acrylic acid, and is recycled to the condensation column via the respective inserted tubes immediately below the second collecting tray, together with the mother liquor (66,487 kg / h) which is heated to 90°C by indirect heat exchange with the withdrawn crude acrylic acid and steam as a heat medium, to the dual flow tray following below the second collecting tray.

[0570] 115000 kg / hr of crude acrylic acid withdrawn from the second collection tray is recycled at a temperature of 96.8° C. via a nozzle directly above the second collection tray spray to keep this section wet and prevent contamination.

[0571] 88650 kg / hour of crude acrylic acid withdrawn from the second collecting tray is cooled by multi-stage indirect heat exchange to a temperature of 29° C. About 1000 kg / hour of deionized water is then added to the cooled acrylic acid.

[0572] The resulting mixture is cooled and subsequently crystallized in a cooled disk crystallizer, and the resulting crystal suspension is purified in a hydraulic melt wash column as described in WO 2008 / 090190, pages 49 and 50. From the melt circuit, stabilized by the addition of a solution of MEHQ in glacial acrylic acid, 22162 kg / h of the desired separated acrylic acid are withdrawn as glacial acrylic acid, which has the following content:

[0573] [Table 69]

[0574] Inhibitor solution 1 is prepared by dissolving 11 kg / h of PTZ (phenothiazine) in 904 kg / h of the abovementioned glacial acrylic acid at 25° C. The remaining stream of glacial acrylic acid, continuously withdrawn from the melting circuit, is continuously fed to a storage tank.

[0575] The mother liquor removed in the wash column is first conducted to a heatable collecting vessel and from there to a tank from which it is heated to 90° C. by thermal integration (as already described) and recycled to the condensation column with a mass flow rate of 66487 kg / h (as a mixed stream (92.2° C., 1.10 bar)) together with 32426 kg / h of crude acrylic acid withdrawn on the second collecting tray of the condensation column in the upper part of the series of dual flow trays below the second collecting tray.

[0576] The composition of the recycled mother liquor is as follows:

[0577] [Table 70]

[0578] 834535 kg / hr of acid water is withdrawn as a second side draw from the third collection tray at a temperature of 53.4° C. and a pressure of 1.229 bar.

[0579] Acid water has the following content:

[0580] [Table 71]

[0581] The withdrawn 27,112 kg / hr of acid water (53.4° C.) is recycled to the top Soman tray together with 35 kg / hr of inhibitor solution 1 (25° C.) and 28 kg / hr of molten MEHQ (T=80° C.). 880 kg / hr of inhibitor solution 1 is recycled to the single-flow Soman tray arrangement (temperature 25° C.) after approximately two-thirds of the length of this Soman separation section (counting from its bottom).

[0582] Pulled out 458.31m 3 / hour of acid water, at a temperature of 28.4°C, is recycled to the centre of an equidistant arrangement of valve trays, which are described in more detail below (cooling is carried out by means of multi-stage indirect heat exchange).

[0583] The withdrawn 325000 kg / h of acid water are recycled at a temperature of 23.4° C. to the top of the above-mentioned equidistant arrangement of valve trays (cooling is carried out together with the above-mentioned amount of acid water by means of multi-stage indirect heat exchange; the final cooling stage from 28.4° C. to 23.4° C. is carried out thermally and with the aid of heat integration (liquid chemical-grade propene is used as coolant and is evaporated simultaneously; the gaseous propene produced is subsequently used to constitute the starting reaction gas mixture 1 for the first stage of the heterogeneously catalyzed gas-phase partial oxidation of propene).

[0584] The withdrawn 7423 kg / h of acid water is fed to an extraction column for the purpose of the subsequent extraction.

[0585] Above the third collecting tray in the condensation column, a series of two-flow valve trays are fixed in an equidistant arrangement. The pressure at the top of the condensation column is 1.17 bar. At the top of the condensation column, 186030 kg / h of residual gas mixture R leaves the separation column at a temperature of 24° C. and with the following content:

[0586] [Table 72]

[0587] In an indirect heat exchanger, the gas mixture R is heated to 32° C. 92920 kg / h thereof are discharged as off-gas stream O from the production process.

[0588] 93110 kg / h of said gas mixture R are compressed by means of a compressor at a pressure of 2.5 bar, raising its temperature to approximately 150° C. 16000 kg / h of it are fed to a first stripping column in the separation zone for stripping the extract from the acid water extraction. The remaining 77110 kg / h of the compressed gas mixture (as cycle gas) are recycled as gas stream P to the first reaction stage 1 as feed stream for preparing the stream of starting reaction gas mixture 1. The sum of the mass flows of the off-gas stream O and the cycle gas stream P constitutes the total mass flow M of the remaining gas mixture R continuously led off from the separation zone. It contains, as pure compounds, 99.99 mol % of the total molar amount of those compounds contained in the product gas mixture 2 continuously transferred to the separation zone, having a boiling point below 250 K at a working pressure of 1 bar.

[0589] The mass flow rate of the recycle gas stream P is 45.35% of the total mass flow rate M of the remaining gas mixture R which is continuously led off from the separation zone.

[0590] 7423 kg / h of acid water to be extracted (temperature=53.4° C.) are fed to the extraction column below the bottom packing via a tubular distributor with suitable passage openings. Above the top packing of the extraction column, a mixture of fresh Palatinol® A and extraction solvent recycled from the first stripping column and previously stripped free therein (ratio of the respective mass flows 1:240) (temperature=50° C.) is introduced at a mass flow rate essentially equal to the 7423 kg / h of acid water. The recycled extraction solvent has the following content:

[0591] [Table 73]

[0592] The specific mass of acid water (at temperature = 53.4°C) is 1018.2 kg / m 3 The extraction solvent is likewise introduced through a tubular distributor with suitable passage openings. The acid water forms a continuous phase, while the extraction solvent forms a dispersed phase in the form of droplets (liquid particle diameter is in the range of 2-5 mm), which descends within the aqueous phase. At the top of the extraction column a raffinate is withdrawn. This raffinate is sent to incineration. The extract is withdrawn from the bottom of the extraction column. It contains the acrylic acid extracted from the acid water. The entire extract is led to the top of the first stripping column. The extract is preheated to 95 ° C by indirect heat exchange in a plate heat exchanger. The heat carrier is the bottom liquid withdrawn from the first stripping column.

[0593] Below the lowest dual flow tray, 16000 kg / h of compressed residue gas mixture R (2.5 bar, 150° C.) is led to a first stripping column, which rises in countercurrent to the descending extract in the stripping column. At the top of the first stripping column, a first contained gas (containing, inter alia, water and acrylic acid) is led off (16875 kg / h) and fed to a second stripping column. The composition of the first contained gas is as follows:

[0594] [Table 74]

[0595] Bottom liquid is continuously withdrawn from the bottom of the first stripping column. A portion of the bottom liquid withdrawn from the first stripping column is cooled in a two-stage indirect heat exchange (the first stage is a plate heat exchanger with thermal integration against the extraction) and recycled to the top of the extraction column as the extractant. Another portion of the bottom liquid withdrawn from the first stripping column is heated to about 160° C. in an external forced circulation tube bundle flash evaporator and recycled to the bottom of the first stripping column.

[0596] The overall propene conversion (based on multiple passes) in this Example 1 is 97.50 mol % (calculated from the difference between the propene content of the chemical-grade propene fed to the production process via the initial reaction gas mixture 1 and the propene content of the off-gas stream O discharged from the production process).

[0597] The total yield of acrylic acid produced in Example 1 is 85.1 mol % (calculated from the acrylic acid as the glacial acrylic acid content continuously fed to the storage tank and the propene as the chemical grade propene content fed to the production process via the initial reaction gas mixture 1).

[0598] From the product gas mixture 2, which is contained in acrylic acid, 99.27 mol % passes into the liquid phase in the separation zone.

[0599] Production of synthesis gas from off-gas in this comparative example The molecular nitrogen content in the off-gas resulting in this comparative example is too high to make economically viable synthesis gas production therefrom. [Explanation of symbols]

[0600] 1 Optional, water vapor flow 2 Optionally, one or more hydrocarbon streams selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane, and ethane. 3. Propene Feed Stream 4. Source flow of molecular oxygen 5 Starting reaction gas mixture 1 6 First reaction step 7 Product Gas Mixture 1 8 Optional, Secondary Gas Stream 9 Starting reaction gas mixture 2 10 Second reaction step 11 Product gas mixture 2 12 Separation Zone 13 A stream of the target product acrylic acid of its desired purity (produced via work-up of the liquid phase 15 in a separation zone), e.g., fed to a storage tank 14 Off-gas stream O discharged from the production process and fed to synthesis gas production 15 liquid phase, the target product acrylic acid migrates from product gas mixture 2 to this liquid phase 16 Residual gas mixture R 17 Recycle gas flow P 18 Optional flow of residual gas mixture R circulating in the separation zone 17+14 Total mass flow rate M of the residual gas mixture R continuously led off from the separation zone 19 High boiling point solvents coming from the separation zone (e.g., separated from the target product acrylic acid via workup of the liquid phase 15 in the separation zone)

Claims

1. 1. A process for the continuous production of either acrolein or acrylic acid as a target product from propene under steady-state operating conditions, comprising: A) As an essential step, in the first reaction step, 2-15% by volume of propene, 2.4 to 37.5% by volume of molecular oxygen, 0.5 to 20% by volume of water, 0-20% by volume of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - O in the range of 1.2 to 2.5 2 Molecular oxygen and propene in the molar ratio: - and CO in the range of 1 to 4 2 Carbon dioxide and carbon monoxide in a molar ratio of CO:CO, A stream of initial reaction gas mixture 1 containing - if the initial reactant gas mixture 1 stream consists of molecular nitrogen, it contains molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5; The flow of the initial reaction gas mixture 1 is subjected to a reaction at a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and an associated selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe), such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2-15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2-30% by volume of water, 0-20% by volume of molecular nitrogen, 0.5 to 30% by volume of carbon monoxide, 1 to 65% by volume of carbon dioxide, 0-80% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - O in the range of 0.5 to 3 2 Molecular oxygen and acrolein in the molar ratio of : acrolein, and - CO in the range of 1 to 4 2 Carbon dioxide and carbon monoxide in the molar ratio of 0.01:0 a stream of the initial reaction gas mixture 2 thus produced, which contains - if the initial reactant gas mixture 2 stream consists of molecular nitrogen, it contains molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.25; The stream of the initial reaction gas mixture 2 is subjected to a single pass acrolein conversion of 95 mol % to 99.99 mol %, at a reaction temperature of 200° C. to 450° C. and an operating pressure of 1 bar to 6 bar, with an associated selectivity S for the formation of the target product acrylic acid, based on the converted propene, evaluated over both reaction stages. AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target products contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. Including, a) the mass flow rate of the recycle gas stream P is at least 65% but not more than 99.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 35% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.01 to 1% by volume of propene, 1-10% by volume of molecular oxygen, 0.5 to 25% by volume of water, 0-20% by volume of molecular nitrogen, 0.5 to 40% by volume of carbon monoxide, 1 to 75% by volume of carbon dioxide, 0-92% by volume of one or more hydrocarbons selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene; and 0.01-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid That is, the method.

2. The first fixed catalyst bed of the first reaction stage comprises a catalyst body, the active material of which is represented by the general formula I Know 12 Like a Want b X 1 c X 2 d X 3 e X 4 f SHE n (I) wherein the variables are defined as follows: X 1 = nickel and / or cobalt, X 2 = thallium, alkali metals and / or alkaline earth metals, X 3 = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten, X 4 = silicon, aluminium, titanium and / or zirconium, a=0.5～5、 b=2～4、 c=3～10、 d=0.02～2、 e=0～5、 f=0 to 10 and n = a number determined by the valence and frequency of the elements other than oxygen in I. The method of claim 1, wherein the at least one multimetal oxide is selected from the group consisting of:

3. The second fixed catalyst bed of the second reaction stage comprises a catalyst body, the active material of which is represented by general formula III With 12 V a X 1 b X 2 c X 3 d X 4 e X 5 f X 6 g OF n (III) wherein the variables are 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=1～6、 b=0.2～4、 c=0.5～18、 d=0～40、 e=0～2、 f=0～4、 g=0 to 40 and n = a number determined by the valence and frequency of the elements other than oxygen in III.

3. The method according to claim 1 or 2, wherein the at least one multimetal oxide is selected from the group consisting of:

4. 4. The process according to claim 1, wherein the hourly space velocity of propene over the first fixed catalyst bed in the first reaction stage is between 50 and 300 l(STP) / l-h.

5. 5. The process according to claim 1, wherein the hourly space velocity of acrolein over the second fixed catalyst bed in the second reaction stage is between 40 and 290 l(STP) / l-h.

6. The starting reaction gas mixture 1 is - recycled gas flow P; - Propene source; - A source of molecular oxygen; - optionally a flow of water vapor; and - optionally one or more streams of one or more hydrocarbons selected from the group consisting of methane, natural gas, ethane, propane, n-butane, isobutane and ethene; The process according to any one of claims 1 to 5, wherein the feed stream comprises

7. The ratio Q in the starting reaction gas mixture 1 = (n CO ) / (n CO +n 炭化水素 7. The method of claim 1, wherein the ratio of β to β is in the range of >0 to 0.

8.

8. The initial reaction gas mixture 1 contains at least 90% by volume O 2 and 10% by volume or less of N 2 8. The process according to claim 1 , wherein the oxygen is formed from a feed stream comprising oxygen of a purity of

9. 9. The method according to claim 1 , wherein the mass flow rate of the recycle gas stream P is at least 70% but not more than 99.5% of the mass flow rate M of the residual gas mixture R continuously led off from the separation zone, and the mass flow rate of the off-gas stream O is not more than 30% but at least 0.5% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone.

10. Under steady-state operating conditions, A) As an essential step, in the first reaction step, 2-15% by volume of propene, 2.4 to 37.5% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0-10% by volume (preferably 0-5%) of molecular nitrogen, 15-25% by volume of carbon monoxide, 20 to 55% by volume of carbon dioxide, 10-30% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially propane); and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - O in the range of 1.2 to 2.5 2 Molecular oxygen and propene in the molar ratio: - and CO in the range of 1 to 4 2 Carbon dioxide and carbon monoxide in a molar ratio of CO:CO, A stream of initial reaction gas mixture 1 containing - if the initial reactant gas mixture 1 stream consists of molecular nitrogen, it contains molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5; The flow of the initial reaction gas mixture 1 is subjected to a reaction at a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and an associated selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe), such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in a second reaction step, 2-15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2-20% by volume of water, 0-10% by volume (preferably 0-5%) of molecular nitrogen, 15-25% by volume of carbon monoxide, 20 to 55% by volume of carbon dioxide, 10-30% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially propane); and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - O in the range of 0.5 to 3 2 Molecular oxygen and acrolein in the molar ratio of : acrolein, and - CO in the range of 1 to 4 2 Carbon dioxide and carbon monoxide in the molar ratio of 0.01:0 a stream of the initial reaction gas mixture 2 thus produced, which contains - if the initial reactant gas mixture 2 stream consists of molecular nitrogen, it contains molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.25; The stream of the initial reaction gas mixture 2 is subjected to a single pass acrolein conversion of 95 mol % to 99.99 mol %, at a reaction temperature of 200° C. to 450° C. and an operating pressure of 1 bar to 6 bar, with an associated selectivity S for the formation of the target product acrylic acid, based on the converted propene, evaluated over both reaction stages. AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target product contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. Including, a) the mass flow rate of the recycle gas stream P is at least 90% but not more than 99% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 10% but at least 1% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.05 to 0.5% by volume of propene, 1-4% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0-10% by volume (preferably 0-5%) of molecular nitrogen, 15-30% by volume of carbon monoxide, 30-60% by volume of carbon dioxide, 10-50% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene (preferably essentially propane); and 0.01-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid The method according to any one of claims 1 to 7, wherein

11. % and ≦99.9 vol.% O 2 and ≦1% by volume and ≧0.1% by volume of N 2 and a recycle gas stream P.

12. Under steady-state operating conditions, A) As an essential step, in the first reaction step, 2-15% by volume of propene, 2.4 to 37.5% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0-10% by volume (preferably 0-5%) of molecular nitrogen, 1.2 to 15% by volume of carbon monoxide, 3 to 30% by volume of carbon dioxide, 40-80% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene, preferably at least 80% by mole of the total molar amount of these hydrocarbons being methane; and >0-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid It consists of: - O in the range of 1.2 to 2.5 2 Molecular oxygen and propene in the molar ratio: - CO in the range of 1 to 4 2 Carbon dioxide and carbon monoxide in a molar ratio of CO:CO, - At least 0.5 O 2 :N 2 Molecular oxygen and molecular nitrogen in a molar ratio of - Carbon monoxide and total molar amounts (quantities) of hydrocarbons with a CO:total hydrocarbons molar ratio in the range of >0 to 0.3 A stream of initial reaction gas mixture 1 containing - if the initial reactant gas mixture 1 stream consists of molecular nitrogen, it contains molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.5; The flow of the initial reaction gas mixture 1 is subjected to a reaction at a reaction temperature of 250° C. to 500° C. and an operating pressure of 1 bar to 6 bar, with a propene conversion in a single pass of 80 mol % to 99 mol % and an associated selectivity S for the formation of the target product acrolein. A obtaining a product gas mixture 1 stream containing the target product acrolein leaving the first reaction stage by continuously passing the product gas mixture 1 through a first fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo), bismuth (Bi) and iron (Fe), such that the acrolein content is between 70 mol % and 99 mol %; and B) If acrylic acid is the target product, Optionally reducing the temperature of the product gas mixture leaving the first reaction stage by direct, indirect or direct and indirect cooling; Optionally, adding a secondary gas to the product gas mixture in the form of molecular oxygen, or an inert gas, or molecular oxygen and an inert gas; and Then, in the second reaction step 2-15% by volume of acrolein, 1 to 45% by volume of molecular oxygen, 2-20% by volume of water, 0-10% by volume (preferably 0-5%) of molecular nitrogen, 1.2 to 15% by volume of carbon monoxide, 3 to 30% by volume of carbon dioxide, 40-80% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene, preferably at least 80% by mole of the total molar amount of these hydrocarbons being methane; and >0-5% by volume of other compounds including, but not limited to, propene and / or acrylic acid It consists of: - O in the range of 0.5 to 3 2 Molecular oxygen and acrolein in the molar ratio of : acrolein, - CO in the range of 1 to 4 2 Carbon dioxide and carbon monoxide in a molar ratio of 0.1:0.05 to 0.1:0.05; and - Carbon monoxide and total molar amounts of hydrocarbons (including acrolein) in a CO:total hydrocarbons molar ratio range of >0 to 0.3 a stream of the initial reaction gas mixture 2 thus produced, which contains - if the initial reactant gas mixture 2 stream consists of molecular nitrogen, it contains molecular oxygen and molecular nitrogen in a molar ratio O2:N2 of at least 0.25; The stream of the initial reaction gas mixture 2 is subjected to a single pass acrolein conversion of 95 mol % to 99.99 mol %, at a reaction temperature of 200° C. to 450° C. and an operating pressure of 1 bar to 6 bar, with an associated selectivity S for the formation of the target product acrylic acid, based on the converted propene, evaluated over both reaction stages. AA obtaining a product gas mixture 2 stream containing the target product acrylic acid leaving the second reaction stage by continuously passing the product gas mixture 2 through a second fixed catalyst bed comprising a shaped catalyst body, the active material of which is at least one multimetal oxide containing molybdenum (Mo) and vanadium (V), such that the product gas mixture 2 stream contains the target product acrylic acid leaving the second reaction stage; and C) continuously transferring either a stream of product gas mixture 1 containing the target product acrolein or a stream of product gas mixture 2 containing the target product acrylic acid to a separation zone, separating the respective target products from the respective product gas mixture in the separation zone by transferring at least 95 mol % of the respective target product contained in the respective product gas mixture from the gas phase to the liquid phase, and continuously withdrawing from the separation zone a total mass flow rate M of a residual gas mixture R which contains, as pure compounds, at least 95 mol % of the total molar amount of compounds contained in the product gas mixtures transferred continuously to the separation zone and which have a boiling point of less than 250 K at an operating pressure of 1 bar, and D) recycling a partial flow of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone into the first reaction stage 1 as a feed stream for preparing a stream of the initial reaction gas mixture 1 containing gas flow P as a gas flow P having the same composition as the residual gas mixture R, and discharging the difference between the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone and the mass flow rate of the recycled gas flow P from the production process as an off-gas stream O having the same composition as the residual gas mixture R. Including, a) the mass flow rate of the recycle gas stream P is at least 75% but not more than 98% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone, b) the mass flow rate of the off-gas stream O is less than or equal to 25% but at least 2% of the total mass flow rate M of the residual gas mixture R continuously led off from the separation zone; c) The composition C of the residual gas mixture R is 0.05 to 0.8 volume percent propene, 1-4% by volume of molecular oxygen, 0.5 to 10% by volume of water, 0-10% by volume (preferably 0-5%) of molecular nitrogen, 1.2 to 15% by volume of carbon monoxide, 3 to 30% by volume of carbon dioxide, 45-90% by volume of at least one hydrocarbon selected from the group consisting of methane, ethane, propane, n-butane, isobutane and ethene, preferably at least 80% by mole of the total molar amount of these hydrocarbons being methane, and 0.01-5% by volume of other compounds including, but not limited to, acrolein and / or acrylic acid The method according to any one of claims 1 to 7, wherein

13. The initial reaction gas mixture 1 is a stream of a propene feedstock of polymer grade (a propene feedstock containing ≥ 99 vol. % propene and ≤ 1 vol. % propane) or chemical grade propene (a propene feedstock containing ≥ 90 vol. % propene and ≤ 10 vol. % but ≥ 2 vol. % propane), ≥ 99.0 vol. % and ≤ 99.9 vol. % O 2 and ≦1% by volume and ≧0.1% by volume of N 2 13. The process according to claim 12, wherein the feed stream is formed from a stream of molecular oxygen having a purity of about 100 to about 1000 nm, a recycle gas stream P, and an additional stream of at least one hydrocarbon selected from the group consisting of methane, natural gas, ethene, propane, n-butane, isobutane and ethene.

14. 14. The method according to any one of claims 1 to 13, wherein transferring each target product from the gas phase to the liquid phase comprises absorbing each target product into a liquid absorbent and / or condensing each target product from the gas phase in at least one separation column containing separating internals through which the respective product gas mixture is directed upwardly from below.

15. 15. A process according to any one of claims 1 to 14, characterized in that it operates in combination with another process and comprises the steps of converting an off-gas stream O discharged from the process and fed to said other process into synthesis gas, and / or mixing said off-gas stream with another synthesis gas produced in said other process from any feedstock other than the off-gas O, to produce the other synthesis gas thus modified.

16. The conversion of the off-gas stream O2 into at least one of the following reactions: - heterogeneously catalyzed steam reforming in the gas phase of the hydrocarbons contained in the off-gas stream O; - CO contained in the off-gas stream O 2 Heterogeneously catalyzed hydrogenation in the gas phase of; - CO, both contained in the off-gas stream O 2 Heterogeneously catalyzed dry reforming in the gas phase of hydrocarbons having - partial gas phase combustion of hydrocarbons and / or compounds composed of C, H and O and contained in the off-gas stream O The method of claim 15, comprising:

17. 17. A method according to claim 15 or 16, wherein the synthesis gas produced and / or the modified other synthesis gas is converted in a synthesis method to produce any chemical compound containing the elements C and H, and optionally O.

Images