DE102022004204A1 - Process for the preparation of maleic anhydride - Google Patents

Abstract

The invention relates to a process for the preparation of maleic anhydride (MA) by catalytic partial gas phase oxidation of saturated and/or unsaturated C4 hydrocarbons, characterized in that the oxidation is carried out in two reactors connected in series, wherein the first reactor is a cooled, predominantly isothermally operated reactor system 1 (main reactor) and in which at least 90% of the total conversion of C4 hydrocarbons takes place. The product gas mixture leaving the first reactor is then passed via a temperature control unit 7 to a second, predominantly adiabatically operated reactor system 8 (post-reactor), which is loaded with a spatially structured VPO catalyst 9 and in which a portion of the C4 hydrocarbons still present in the product gas (residual C4-KW) are converted to further MA and a portion of the by-products formed in the main reactor are broken down, under the proviso that the maximum temperature of the reaction gas occurring in the post-reactor 8 is always below the maximum temperature occurring in the main reactor.

Description

With an annual worldwide production of well over 2 million tonnes, maleic anhydride (MA) is an important intermediate product for the production of polymers, solvents, dyes, etc. ( F. Trifiro, RK Grasselli, Topics in Catalysis 57 (14), 1188 - 1195 ). The production of MSA is carried out for the most part via a gas phase oxidation of saturated and unsaturated C4 hydrocarbons with oxygen or oxygen-containing gas mixtures, for example air, and has been known for many years. The oxidation can take place in various reactor types (fixed bed, fluidized bed or transport bed reactor) (Ullmann, Encyclopedia of Industrial Chemistry, 2017, electronic edition, chapter Maleic and Fumaric Acid). The most commonly used reactor type for MSA synthesis is a salt bath cooled fixed bed reactor, consisting of a large number of reaction tubes (tube bundle reactor) filled with a bed of shaped catalyst bodies. The catalysts used are usually full catalysts based on vanadyl pyrophosphate (VO) ₂ P ₂ O ₇ , the so-called VPO catalysts.

The isolation of the MSA is carried out by cooling the reaction gas mixture leaving the reactor and a subsequent gas scrubbing, usually with an organic solvent that is characterized by good solubility for maleic anhydride (e.g. dibutyl phthalate, WO 2012/081043 A1 ) or with water to form maleic acid, followed by further sub-steps such as concentration, dewatering, dehydration and distillation. In the distillation step, the by-products formed during the oxidation reaction, mainly acetic and acrylic acid, are also separated. Acrylic acid is particularly undesirable because it can lead to polymer deposits and thus problems in the processing steps and also has a highly negative impact on the heat stability of the pure MA, an important quality feature. The remaining reaction gas freed from the target product MA, mainly consisting of CO/CO ₂ , residual oxygen and inert gas (e.g. nitrogen) as well as unreacted C4 hydrocarbons, is fed into exhaust gas combustion for thermal utilization in most current MA production plants.

The production of the VPO catalyst mass is usually carried out in 3 steps, the synthesis of the precursor compound VO(HPO4)*0.5 H2O, the shaping and a calcination step to convert the precursor compound into the actual catalytically active VPO compound. All of these steps influence the final performance of the catalyst. To control the catalytic properties, most VPO catalysts are additionally doped with promoter compounds, which can in principle be added at any time during production. Recent studies also demonstrate the important role of porosity, micropores and macropores on the activity and selectivity of the MSA catalyst ( Y. Dong. FJ Keil, O. Korup, F. Rosowski, R.Horn, Chem. Eng. Science 142 (2016) pp. 299 - 309 )

The synthesis of the precursor compound is preferably carried out by the so-called organic route by reacting V ₂ O ₅ with a reducing agent and H ₃ PO ₄ in an organic solvent such as i-butanol, the latter simultaneously serving as a reducing agent ( US4132670 ). The addition of H ₃ PO ₃ as an additional reducing agent and P source is also described ( US4382876 ). Furthermore, during precursor production, entrainers can be used to facilitate the removal of water, such as cyclohexane and so-called structure formers from the groups of alcohols, organic amine or phosphorus compounds ( EP1117482B1 ) may be added.

The shaping is usually carried out by an extrusion process or tableting of the dried precursor compound. The dimensions of the molded bodies are adapted to the geometry of the reaction tubes in the fixed bed reactor in order to minimize or optimize the pressure loss that occurs. Typical tube diameters of the reaction tubes arranged in the fixed bed reactors are 21 - 25 mm, the tubes are usually 3 - 6 m long. The molded bodies are usually so-called full contacts, ie they consist almost exclusively of catalyst mass, and are in the form of cylinders, hollow cylinders ( US$4,283,307 ; WO 01/68245 ), Trilobes ( EP0593646 ) and other special shapes such as hollow cylinders with four corners ( WO 2012/069481 A1 ) or hollow cylinders with two or more through internal bores ( WO 2010/072723 A2 ) are used. It is also common to add extrusion or tableting aids and pore formers to the catalyst mass before the shaping step. The dimensions of the cylindrical molded bodies are between 3 and 10 mm high and 3 - 10 mm in diameter. The diameters of the internal cylinder cavities (inner bores) are between 0.5 and 5 mm.

The third step, the calcination of the dried precursor compound to the actual active VPO catalyst mass, the vanadyl pyrophosphate, is the elimination of the water of hydration. The calcination can be carried out discontinuously, for example in a shaft furnace or tray furnace, as well as continuously in a rotary tube or belt calciner or directly in the MSA reactor. The processes described in the patent literature differ in terms of heating rate, composition of the gas atmosphere, holding times and cooling rates. The aim is to adjust the calcination conditions to the preferred degree of oxidation of the vanadium for the catalyst. Values of + 4.0 to +4.3 are usually set ( EP1117482B1 , US5137860 , WO2010/072723 A2 ).

Also described are so-called MSA shell catalysts, which are produced by coating inert support materials with the precursor compound, the hemihydrate VO(HPO ₄ )*0.5H ₂ O ( EP0756519B1 ), or after the EP0917909B1 by coating an inert carrier material with a VPO mass that is already catalytically active. In both variants, the geometry of the catalyst molding is determined by the inert carrier material.

To control the catalytic properties, most VPO catalysts are additionally doped with promoter compounds. In principle, these can be added at any time during production. This usually takes place during the wet preparation of the precursor compound. They are introduced in the form of soluble promoter compounds such as acetates, acetylacetonates, oxalates, oxides, etc. Promoter elements include Co, Mo, Zn, Fe, Li, Ni, Zr, Ce, Bi, Cr. The V/promoter element ratios are between 1:0.2 and 0.001 ( US4132670 , EP0458541 ).

The reaction of C4 hydrocarbons with air or oxygen-containing gas mixtures in the fixed bed reactor is carried out at temperatures between 380 and 430 °C, depending on the catalyst used. The usual gas velocities (GHSV) are between 1500 and 2000 h ^-1 , the C4 hydrocarbon concentrations are between 1.5 and 2.2 vol.%. For reasons of catalyst stability with regard to performance and service life, it is state of the art that volatile P compounds such as trimethyl phosphate in the range of 2 - 20 ppm and water vapor are added to the reactant gas under the reaction conditions ( US5185455 ). Due to the high heat of the target reaction and the not insignificant side reaction, the total oxidation to CO/CO ₂ , a temperature profile is formed in the catalyst bed despite the heat removal through the salt bath, with the temperature maximum (hot spot) occurring within the first third of the catalyst bed. Depending on the reaction conditions and catalyst performance, the hot spot temperatures can be between 20 and 60 °C above the corresponding salt bath temperature, i.e. temperatures between 400 and 460 °C can be reached. Temperatures above 440 °C lead to a significant reduction in MSA selectivity (Carsten Becker, (2002) Dissertation: Katalytische Wandreaktorkonzepte für MSA-Synthese und Methanol-Dampfreformierung, Chapter 8, University of Stuttgart). Depending on the activity of the catalyst, the reaction can sometimes run away. There is therefore a maximum temperature depending on the catalyst that should not be exceeded.

In the multi-tube reactors commonly used, temperature sensitivity limits the conversion of the C4-HC used to around 85% and thus also the economic viability of this process. This is a limitation that has only been eliminated to a limited extent despite the development and introduction of new, improved VPO catalysts. There has therefore been no lack of attempts in the past to reduce the significantly high C4 loss by changing the process. A reduction in the particularly troublesome by-product acrylic acid can also contribute to improving the economic viability of the MSA process.

State of the art

In the US3904652 It is proposed to return the C4-containing scrubber exhaust gas to the reactor in a kind of cycle process and thus to return the unused C4-KW portion to the gas phase oxidation. In the US6194587B1 A process is claimed in which at least 2 isothermally operated fixed bed reactors are connected in series, whereby the MA formed in the first reactor is largely removed or isolated from the reaction gas via a gas scrubber. The gas mixture leaving the gas scrubber is then again loaded with C4 feed and possibly oxygen and then converted to MA again in the second reactor. Such a cascade can also consist of three or more reactors.

Another process variant consists in the separation of the cooling circuit in the fixed bed reactor into two or more cooling segments, a so-called 2- or multi-zone reactor, whereby the reaction control of the MSA synthesis with two or more separate cooling circuits is intended to achieve an increase in yield compared to the standard procedure ( WO 01/68626 A1 , EP3360611A1 ). This reaction principle can be further optimized by using a VPO catalyst optimized for the reaction process (in form and composition) in each zone. Patents describe the use of two-layer VPO catalyst systems in which the 2nd catalyst layer with tungsten ( EP3771490B1 ) or molybdenum ( US5929256 ). Such systems are characterized by the fact that the formation of the disturbing by-product, acrylic acid, can be reduced. In the US6120654 An optimized distillation variant is described which improves the separation of acrylic acid, thereby achieving a significant increase in the color stability of the MSA.

Task

Ultimately, none of these process variants have been able to prevail. All of the above proposals require high additional investments (e.g. more complex reactors, additional isothermal reactors and larger fans) and a considerably higher energy requirement (heating energy, higher pressure losses). Furthermore, existing MSA systems can only be converted at great expense.

• - eine höhere MSA-Ausbeute durch einen gesteigerten C4-Umsatz realisieren lässt,
• - die Konzentration an unerwünschten Nebenprodukten, insbesonders die Acrylsäure reduziert werden kann,
• - größere Investitionen, z.B. in Form eines zusätzlichen Salzbadreaktors, vermieden werden, und
• - sich in bereits vorhandene MSA-Anlagen integrieren lässt.

The task was therefore to find a process variant with which

• - a higher MSA yield can be achieved through increased C4 conversion,
• - the concentration of undesirable by-products, especially acrylic acid, can be reduced,
• - major investments, e.g. in the form of an additional salt bath reactor, can be avoided, and
• - can be integrated into existing MSA systems.

Subject of the invention

1. a) dem Einbau mindestens einer zusätzlichen, vorwiegend adiabatisch betriebenen Reaktionseinheit (8) zwischen der vorwiegend isotherm betriebenen Reaktionseinheit 1 (Hauptreaktor) und dem Aufarbeitungsteil 11, wobei der Hauptumsatz an C4-KW in der isotherm betriebenen Reaktionseinheit 1 erfolgt und mindestens 90 % des Gesamtumsatzes, bevorzugt mindestens 95 %, besonders bevorzugt mindestens 98 % betragen soll und der vorwiegend adiabatisch betriebenen Nachreaktor 8 maximal 10 %, bevorzugt maximal 5 %, besonders bevorzugt maximal 2 % zum Gesamt C4-Umsatz beitragen soll, wobei
2. b) der Umsatz an C4-KW im Nachreaktor 8 durch die Temperierung des den Hauptreaktor verlassenden Gesamtgastroms in einer dem Nachreaktor 8 vorgeschalteten Temperiereinheit 7 in der Weise gesteuert wird, dass die im Nachreaktor 8 auftretende Maximaltemperatur des Reaktionsgases nicht höher als die im Hauptreaktor maximal auftretenden Gastemperatur (Hot-Spot Temperatur) ist, und
3. c) dass der im Nachreaktor 8 angeordnete VPO-Katalysator 10 in seiner Gesamtheit eine räumlich strukturierte Geometrie aufweist.

The solution to the problem is in 1 and consists of

1. a) the installation of at least one additional, predominantly adiabatically operated reaction unit (8) between the predominantly isothermally operated reaction unit 1 (main reactor) and the processing section 11, wherein the main conversion of C4-KW takes place in the isothermally operated reaction unit 1 and should amount to at least 90% of the total conversion, preferably at least 95%, particularly preferably at least 98%, and the predominantly adiabatically operated post-reactor 8 should contribute a maximum of 10%, preferably a maximum of 5%, particularly preferably a maximum of 2% to the total C4 conversion, wherein
2. b) the conversion of C4-KW in the post-reactor 8 is controlled by the temperature control of the total gas flow leaving the main reactor in a temperature control unit 7 upstream of the post-reactor 8 in such a way that the maximum temperature of the reaction gas occurring in the post-reactor 8 is not higher than the maximum gas temperature occurring in the main reactor (hot spot temperature), and
3. c) that the VPO catalyst 10 arranged in the post-reactor 8 has a spatially structured geometry as a whole.

The invention will be explained in more detail in the following example. The example does not represent a limitation in the sense of the present invention. The invention can include any embodiment familiar to the person skilled in the art.

Example: Implementation of the method according to the invention

2 shows a schematic of the process principle of the process according to the invention. The reactant gas, consisting of C4-KW/air, is then passed through a tube bundle reactor 2 cooled with a salt bath cooler 4 and operated predominantly isothermally at temperatures between 380 and 420° C. The reaction tubes 3 have a diameter of 21 - 25 mm and a length of approx. 6 m, and are filled with a particulate, ring-shaped VPO catalyst produced and doped via the organic route. The concentration of C4-KW in the reactant gas is generally between 1.4 and 2 vol.%, the gas velocities (GHSV) between 1400 and 2000 h ^-1 . The reactant gas is conditioned to a constant moisture content (2 - 3 vol.% water) before entering the main reactor, and a phosphorus compound such as trimethyl phosphate (2 - 20 ppm) is added after the start-up process.

Due to the high heat of the target reaction and the not insignificant side reaction, the total oxidation to CO/CO ₂ , a temperature profile is formed within the first third of the catalyst bed despite the heat dissipation through the salt bath. The height and position of the temperature profile, as well as the maximum temperature in the catalyst bed (hot spot) are recorded by several multi-thermocouples 5 evenly distributed over the reactor cross-section, which are positioned centrally in appropriately selected reaction tubes, and are used to control the reactor (C4-CHC concentration/conversion/temperature). In this way, it is ensured that the maximum permitted difference between the temperature maximum in the catalyst bed (T _{hot spot} ) and the salt bath temperature (T _{salt bath} , measured at the inflow of the molten salt 6 into the reactor) is less than 60° C.

The product gas leaving the main reactor, mainly consisting of O ₂ /N ₂ /CO/ CO ₂ /MSA/H ₂ O and remaining unreacted C4-KW, is now passed through a downstream, predominantly adiabatically operated reactor unit, the post-reactor 8, instead of to the further processing process 11. The post-reactor is preceded by a separate gas tempering unit 7, e.g. in the form of a finned tube heat exchanger, in which the product gas from the main reactor can be cooled/tempered before entering the actual post-reactor. The adiabatically operated post-reactor is loaded with honeycomb-shaped, Mo-doped VPO catalysts 10, which are produced by coating ceramic cordierite honeycombs with a VPO catalyst material produced by the organic route, such as in EP0917909B1 described, can be produced via a dipping process. The amount of active VPO catalyst mass per honeycomb can be between 50 and 400 g/l honeycomb depending on the desired activity, 80 - 150 g/l honeycomb is preferred. The cell densities (cells per square inch, cpsi) of the honeycombs are between 50 and 400 cpsi, cell densities of 100 - 200 cpsi are preferred. The honeycomb-shaped VPO catalysts are arranged in layers in the post-reactor, usually 1 - 4 layers, preferably 2 - 3 layers, so that the product gas is distributed as evenly as possible over the entire inflow area. It is advantageous to maintain a distance between the layers that allows the gas flow to be mixed before it enters the subsequent catalyst layer. The total amount of VPO catalyst honeycombs is selected so that a space velocity of 10,000 to 30,000 h ^-1 results, with space velocities of 15,000 - 20,000 h ^-1 being preferred.

The catalyst configuration such as catalyst quantity, number of catalyst layers, cell densities, etc. are selected so that the additional pressure loss in the post-reactor can be covered by the existing blower and the conditions in the main reactor are not negatively affected. The post-reactor is operated in such a way that the proportion of C4 conversion achieved in the post-reactor corresponds to a maximum of 10% of the total conversion, preferably a maximum of 5%, particularly preferably a maximum of 2% of the total conversion. The conversion in the post-reactor is controlled by tempering/cooling the product gas leaving the main reactor in the gas cooler unit 7. It should be noted that the maximum gas temperatures occurring in the post-reactor, recorded with representatively distributed thermocouples 9, are not higher than the maximum gas temperatures occurring in the main reactor (hot spot). At the beginning of the start-up phase of the post-reactor, it is advisable to cool the product gas flow more strongly in order to avoid undesirable over-oxidation of the target product (to CO ₂ ) and the associated exceeding of the temperature limits. By slowly increasing the T, the desired additional conversion of the remaining C4-KW portion to MA and the degradation of undesirable by-products such as acrylic acid can be achieved while maintaining a stable operating state. The further processing/isolation of the MA is carried out according to the state of the art by cooling the reaction gas mixture leaving the post-reactor with the usual sub-steps, depending on the design of the plant, such as gas washing, concentration, dewatering, dehydration and distillation as well as thermal utilization of the exhaust gas in the processing section 11.

The process according to the invention leads to a noticeable increase in profitability even with a small improvement in MSA yield of approx. 1% and it can be integrated into existing MSA plants without the need for major investment. For a plant size of 50,000 tons per year, which is typical in industry, an additional yield of 1% MSA already corresponds to approx. 500 t of target product per year. The reaction heat expected from the additional conversion of around 1% C4-KW should, depending on the selectivity of the monolithic VPO catalyst, lead to an increase in the temperature of the reaction gas of around 10 - 15 °C and should therefore also be easy to control in the adiabatically operated post-reactor. By using specifically doped VPO catalysts in the post-reactor, undesirable by-products such as acrylic acid can be significantly reduced, which minimizes the distillation effort and thus also improves profitability.

List of reference symbols

1

Mainly isothermal reaction unit

2

Tube bundle reactor

3

Reaction tubes filled with particulate VPO catalyst

4

Salt bath cooler with pump

5

Reaction tube with centered thermocouple

6

Inflow of the cooled molten salt to the tube bundle reactor

7

Temperature control unit (gas cooler)

8th

predominantly adiabatically operated reactor unit (post-reactor)

9

Thermocouple post-reactor

10

spatially structured VPO catalyst (honeycomb)

11

Reprocessing part

QUOTES INCLUDED IN THE DESCRIPTION

This list of documents listed by the applicant was generated automatically and is included solely for the better information of the reader. The list is not part of the German patent or utility model application. The DPMA accepts no liability for any errors or omissions.

Cited patent literature

• WO 2012081043 A1 [0002]
• US 4132670 [0004, 0008]
• US 4382876 [0004]
• EP 1117482 B1 [0004, 0006]
• US 4283307 [0005]
• WO 0168245 [0005]
• EP0593646 [0005]
• WO 2012069481 A1 [0005]
• WO 2010072723 A2 [0005, 0006]
• US5137860 [0006]
• EP 0756519 B1 [0007]
• EP 0917909 B1 [0007, 0019]
• EP0458541 [0008]
• US5185455 [0009]
• US3904652 [0011]
• US 6194587 B1 [0011]
• WO 0168626 A1 [0012]
• EP 3360611 A1 [0012]
• EP 3771490 B1 [0012]
• US5929256 [0012]
• US6120654 [0012]

Cited non-patent literature

• F. Trifiro, R.K. Grasselli, Topics in Catalysis 57 (14), 1188 - 1195 [0001]
• Y. Dong. F.J. Keil, O. Korup, F. Rosowski, R.Horn, Chem. Eng. Science 142 (2016) pp. 299 - 309 [0003]

Claims (10)

The invention relates to a process for producing maleic anhydride (MA) by a catalytic partial gas phase oxidation reaction of saturated and/or unsaturated C4 hydrocarbons, characterized in that the gas phase oxidation is carried out in a combination of a predominantly isothermal reactor unit 1, the main reactor, and a downstream predominantly adiabatically operated reactor unit 8, the post-reactor, wherein at least 90% of the total C4 conversion takes place in the main reactor and the predominantly adiabatically operated post-reactor is loaded with a spatially structured VPO catalyst 9, in which a partial further conversion of the C4 hydrocarbons (residual C4-HC) still contained in the gas stream from the main reactor to further MA takes place. Process for the production of MSA according to Claim 1 , characterized in that the proportion of total C4-KW conversion in the predominantly adiabatically operated post-reactor 8 is a maximum of 10%, preferably a maximum of 5%, particularly preferably a maximum of 2%. Process for the production of MSA according to Claims 1 and 2 , characterized in that a tempering unit 7 is interposed between the predominantly isothermal reactor unit 1 and the adiabatically operated post-reactor 8, with which the inlet temperature of the gas flow into the post-reactor is regulated in such a way that the maximum reaction temperature occurring in the post-reactor is always lower than the maximum gas temperature (hot spot temperature) occurring in the predominantly isothermal reactor unit 1. Process for the production of MSA according to Claims 1 - 3 , characterized in that the temperature control unit 7 and the predominantly adiabatically operated post-reactor 8 are combined in a common apparatus. Process for the production of MSA according to Claims 1 - 4 , characterized in that the predominantly isothermal reaction unit 1 is a tube bundle reactor 2 provided with a salt bath cooler 4, the tubes 3 of which are filled with particulate VPO catalysts and the VPO catalyst mass contains one or more dopants from the group Li, Fe, Mo, W, Cr, Co, Ni, Ce, Zr, Zn, Bi. Process for the production of MSA according to Claims 1 - 5 , characterized in that the spatially structured VPO catalyst 10 used in the adiabatically guided post-reactor is a monolithic VPO supported catalyst consisting of a monolithic support material, preferably in honeycomb form, which is coated with doped VPO catalyst mass. Process for the production of MSA according to Claims 1 - 6 , characterized in that the support material of the VPO supported catalyst 10 used in the post-reactor consists of one or more materials from the group comprising cordierite, silicates, silicon dioxide, silicon carbide, aluminum oxide, aluminates, metals or metal alloys. Process for the production of MSA according to Claims 1 - 7 , characterized in that the VPO supported catalysts 10 arranged in the adiabatically guided post-reactor have a VPO catalyst mass of 50 - 200 g/l catalyst support. Process for the production of MSA according to the Claims 1 - 8th , characterized in that in the adiabatically controlled post-reactor 8 monolithic VPO supported catalysts 10 with a cell density of 100 - 400 cpsi (cells per square inch) are used and are arranged in one or more layers. Process for the production of MSA according to the Claims 1 - 9 , characterized in that the gas phase oxidation of saturated and/or unsaturated C4 hydrocarbons is carried out in a combination of an isothermally controlled tube bundle reactor 2 filled with particulate VPO catalyst and an adiabatically operated post-reactor 8 loaded with VPO honeycomb catalysts as well as an intermediate gas cooling system 7, wherein the active mass of the VPO honeycomb catalysts 10 arranged in one or more layers contains one or more dopants from the group Li, Fe, Mo, W, Cr, Co, Ni, Ce, Zr, Zn, Bi, preferably one or more from the group Mo, W, Cr, and the composition of the catalytically active mass is the same or different from layer to layer.

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