EP1651588A1 - Method for dehydrogenation of carbonyl compounds - Google Patents

Abstract

A method for the production of alpha, beta unsaturated acyclic or cyclic carbonyl compounds by hydrogenation of corresponding saturated carbonyl compounds in a gaseous phase on a heterogeneous dehydrogenation catalyst, characterized in that the dehydrogenation catalyst contains platinum and/or palladium and tin on an oxidic carrier.

Description

 Nerfahren for dehydrogenation of carbonyl compounds

The invention relates to a ner process for the production of alpha, beta-unsaturated acyclic or cyclic carbonyl compounds by dehydrogenation of the corresponding saturated carbonyl compounds in the gas phase over a heterogeneous dehydrogenation catalyst.

JP 49127909 A2 describes a ner process for the dehydrogenation of saturated ketones. Accordingly, butanone can be converted to l-buten-3-one at 500 ° C. on a catalyst containing iron oxide, aluminum oxide and potassium oxide with the addition of water vapor with 5.5% conversion and 83% selectivity.

However, such catalysts based on iron oxide with a relatively large amount of water vapor in relation to the hydrocarbon to be dehydrogenated must be used and nevertheless deactivate very quickly under the influence of hydrocarbons with oxo functions.

No. 6,433,229 B1 describes a ner process for the preparation of cyclic alpha, beta-unsaturated ketones by dehydrogenation of the corresponding saturated ketones, in particular the dehydrogenation of cyclopentanone to 2-cyclopenten-l-one, in the gas phase over heterogeneous catalysts in the presence of less than 0.1 moles of oxygen per mole of the starting material to be dehydrogenated. CuO, AgO, PdO, O, Mn ₂ O ₃ or Re ₂ O on ZnO, CaO, BaO, SiO ₂ or Al ₂ O _{3 are} described as catalysts. In one example, a catalyst containing 9.5% Pd and 0.55 Pt on ZrO _{2 is} used. The regeneration of the catalysts is said to take place at temperatures between 400 and 500 ° C.

The catalysts described in US Pat. No. 6,433,229 B1 are also deactivated within hours under the conditions described there. In addition, regeneration at temperatures between 400 and 500 ° C is not very practical from a reaction point of view, since in order to avoid higher hot spot temperatures, on the one hand a very strong ner thinning of the oxygen-containing gas with inert gas is required, which makes regeneration long, tedious and expensive. Beyond that Temperatures between 400 and 500 ° C are generally not sufficient to completely remove carbon-containing deposits from the catalysts, in particular from deeper pore systems, within a reasonable regeneration time. On the other hand, however, higher hot spot temperatures can damage the catalysts described, for example by sintering the oxidic materials.

It is therefore an object of the invention to find an improved ner process for the dehydrogenation of hydrocarbons containing oxo functions over heterogeneous catalysts in the gas phase. In particular, this should also include simple and practical regeneration of the catalysts.

The object is achieved by a ner process for the production of alpha, beta-unsaturated acyclic or cyclic carbonyl compounds by dehydrogenation of the corresponding saturated carbonyl compounds in the gas phase on a heterogeneous dehydrogenation catalyst, characterized in that the dehydrogenation catalyst is platinum and / or palladium and tin on an oxidic support contains.

The dehydrogenation catalysts used according to the invention generally have a support and an active composition. The carrier consists of a heat-resistant oxide or mixed oxide. The dehydrogenation catalyst preferably contains a metal oxide as carrier which is selected from the group consisting of zirconium dioxide, zinc oxide, aluminum oxide, silicon dioxide, titanium dioxide, magnesium oxide, lanthanum oxide, cerium oxide and mixtures thereof. Preferred carriers are zirconium dioxide and / or silicon dioxide; mixtures of zirconium dioxide and silicon dioxide are particularly preferred.

The active composition of the dehydrogenation catalyst used according to the invention contains platinum and / or palladium as active metals. The dehydrogenation catalyst also contains tin. In general, the dehydrogenation catalyst contains 0.01 to 2% by weight, preferably 0.1 to 1, particularly preferably 0.2 to 0.6% by weight of palladium and / or platinum and 0.01 to 10% by weight , preferably 0.2 to 2% by weight, particularly preferably 0.4 to 1% by weight of tin, based on the total weight of the dehydrogenation catalyst. If the dehydrogenation catalyst contains platinum as the active metal, the weight ratio of platinum: tin is preferably 1 to 3, in particular approximately 2.

In addition, the dehydrogenation catalyst can have one or more elements of the 1st and / or 2nd main group, preferably potassium and / or cesium. Furthermore, the dehydrogenation catalyst can include one or more elements of III. Νebengruppe including the lanthanides and actinides, preferably lanthanum and / or cerium. Finally, the dehydrogenation catalyst can include one or more elements of III. and / or other elements of the IN. Have main group, preferably one or more elements from the group consisting of boron, gallium, silicon and lead.

In a preferred embodiment, the dehydrogenation catalyst contains, in addition to palladium and / or platinum and tin, at least one element from main group I and / or II and at least one element from III. Subgroup including the lanthanides and actinides.

In advantageous embodiments, the active composition contains the following further components:

- At least one element from main group I or II, preferably cesium and / or potassium, with a content of 0 to 20% by weight, preferably 0.1 to 10% by weight, particularly preferably 0.2 to 1.0% by weight, based on the total weight of the catalyst;

- at least one element of III. Secondary group including the lanthanides and actinides, preferably lanthanum and / or cerium with a content of 0 to 20% by weight, preferably 0.1 to 15% by weight, particularly preferably 0.2 to 10% by weight, in particular from 1 to 5% by weight, based on the total weight of the catalyst.

The dehydrogenation catalyst is preferably halogen-free.

Preparation of the dehydrogenation catalyst Precursors for oxides of zircon, silicon, aluminum, titanium, magnesium, lanthanum or cerium, which can be converted into the oxides by calcination, can be used to produce the dehydrogenation catalysts used according to the invention. These can be prepared by known ner processes, for example by the sol-gel process, precipitation of the salts, dewatering of the corresponding acids, dry mixing, slurrying or spray drying. For example, to produce a ZrO ₂ ^β Al ₂ O ₃ ^* * SiO ₂ mixed oxide, a water-rich zirconium oxide of the general formula ZrO ₂ ^β xH ₂ O can first be produced by precipitation of a suitable precursor containing zirconium. Suitable zirconium precursors are, for example, Zr (,O ₃ ), ZrOCl ₂ , or ZrCl ₄ . The precipitation itself is carried out by adding a base such as NaOH, KOH, Na ₂ CO ₃ and NH ₃ and is described, for example, in EP-A 0 849 224. To produce a ZrO ₂ * SiO ₂ mixed oxide, the zirconium-containing precursor obtained beforehand can be mixed with a silicon-containing precursor. Well suited Precursors for SiO ₂ are e.g. B. water-containing brine of SiO ₂ such as Ludox. The two components can be mixed, for example, by simple mechanical mixing or by spray drying in a spray.

To produce a ZrO ₂ ^* * SiO ₂ ^* Αl ₂ θ ₃ mixed oxide, the SiO ₂ 'ZrO ₂ powder mixture obtained as described above can be mixed with an aluminum-containing precursor. This can be done, for example, by simple mechanical mixing in a kneader. The preparation of the ^β ZrO ₂ SiO ₂ ^«* Al ₂ O ₃ ~ mixed oxide but may also be effected in a single step by dry mixing of the individual precursors.

The supports for the dehydrogenation catalysts used according to the invention have the advantage, among other things, that they can be easily deformed. For this purpose, the powder mixture obtained is mixed with a concentrated acid in the kneader and then in a shaped body, e.g. by means of an extruder or an extruder.

In particular embodiments, the dehydrogenation catalysts used according to the invention have a defined pore structure. When using mixed oxides, there is the possibility of deliberately influencing the pore structure. The grain size of the different precursors influence the pore structure. For example, the use of Al ₂ O ₃ with low loss on ignition and a defined grain size composition allows macropores to be created in the structure. The use of Al ₂ O ₃ with a loss on ignition of about 3% (eg Puralox®) has proven itself in this connection.

A further possibility for the targeted production of the supports with special pore radius distributions for the dehydrogenation catalysts used according to the invention consists in the addition of various polymers during production which are partially or completely removed by calcination, pores being formed in defined pore radius ranges. The polymers and the oxide precursors can be mixed, for example, by simple mechanical mixing or by spray drying in a spray tower.

The use of PVP (polyvinylpyrrolidone) has proven particularly useful for the production of supports with a bimodal pore radius distribution. If this is added to one or more oxide precursors for oxides of the elements Zr, Ti, Al or Si in one production step, macropores in the range from 200 to 5000 nm are formed after the calcination. Another advantage of using PVP is the easier deformability of the carrier. So strands with good mechanical properties can be easily produced from freshly precipitated water-containing ZrO ₂ ^» x H ₂ O, which has previously been dried at 120 ° C, with the addition of PVP and formic acid, without additional oxide precursors.

The calcination of the supports for the dehydrogenation catalysts used according to the invention is advantageously carried out after the application of the active components and is carried out at temperatures from 400 to 1000 ° C., preferably from 500 to 700 ° C., particularly preferably at 550 to 650 ° C. and in particular at 560 to 620 ° C performed. The calcination temperature should usually be at least as high as the reaction temperature of the dehydrogenation at which the dehydrogenation catalysts according to the invention are used.

The supports of the dehydrogenation catalysts used according to the invention generally have high BET surface areas after the calcination. The BET surface areas are generally greater than 40 m Ig, preferably greater than 50 m Ig, particularly preferably greater than 70 m ² / g. The pore volume of the dehydrogenation catalysts according to the invention is usually 0.2 to 0.6 ml / g, preferably 0.25 to 0.5 ml / g. The average pore diameter of the dehydrogenation catalysts according to the invention, which can be determined by mercury porosimetry, is between 3 and 20 nm, preferably between 4 and 15 nm.

A bimodal pore radius distribution is also characteristic of the dehydrogenation catalysts used according to the invention. The pores are in the range up to 20 nm and between 40 and 5000 nm. Based on the total pore volume of the dehydrogenation catalyst, these pores have a total of at least 70%. The proportion of pores smaller than 20 nm is generally between 20 and 60%, the proportion of pores between 40 and 5000 nm is also generally 20 to 60%.

The dehydrogenation-active component is generally applied by impregnation with a suitable metal salt precursor. Instead of impregnation, the dehydrogenation-active component can also be carried out by other methods such as, for example, spraying on the metal salt precursor. Suitable metal salt precursors are, for example, the nitrates, acetates and chlorides of the corresponding metals, and complex anions of the metals used are also possible. Platinum is preferably used as H ₂ PtCl ₆ , platm (II) oxalate or Pt (NO ₃ ) ₂ and palladium as palladium (II) oxalate or Pd (NO ₃ ) ₂ . Water is just as suitable as a solvent for the metal salt precursors, as is organic Solvent. Water and lower alcohols such as methanol and ethanol are particularly suitable.

Suitable precursors when noble metals are used as the dehydrogenation-active component are also the corresponding noble metal sols, which can be prepared by one of the known processes, for example by reducing a metal salt in the presence of a stabilizer such as PVP with a reducing agent. The manufacturing technology is dealt with in detail in German patent application DE 195 00 366.

The further components of the active composition can either be used during the preparation of the carrier, e.g. by common precipitation, or subsequently, for example by impregnating the carrier with suitable precursor compounds. As precursor compounds, use is generally made of compounds which can be converted into the corresponding oxides by calcining. For example, hydroxides, carbonates, oxalates, acetates, chlorides or mixed hydroxycarbonates of the corresponding metals are suitable.

To apply alkali and alkaline earth metals, aqueous solutions of compounds which can be converted into the corresponding oxides by calcination are expediently used. Hydroxides, carbonates, oxalates, acetates or basic carbonates of the alkali or alkaline earth metals are suitable, for example. If the catalyst carrier with metals of III. Main or sub-group doped, the hydroxides, carbonates, nitrates, acetates, formates or oxalates are often used, which can be converted into the corresponding oxides by calcination, for example La (OH) ₃ , La ₃ (CO ₃ ) ₂ , La (NO ₃ ) ₃ , lanthanum acetate, lanthanum formate or lanthanum oxalate.

The dehydrogenation catalyst can be fixed in the reactor or e.g. be used in the form of a fluidized bed and have a corresponding shape. Suitable are e.g. Forms such as grit, tablets, monoliths, spheres or extrudates (strands, wagon wheels, stars, rings).

Preparation of a preferred catalyst support

In a particularly preferred embodiment of the process according to the invention, a catalyst support is used which can be obtained by mixing zirconium dioxide powder with a monomeric, oligomeric or polymeric organosilicon compound as

Binder, optionally a pore former, optionally an acid, water and optionally further additives are mixed to form a kneadable mass, the mass is homogenized, shaped into shaped bodies, dried and calcined.

By mixing essentially monoclinic zirconium dioxide powder, which has a high surface area, with an organosilicon compound as binder, which forms SiO ₂ when calcined, shaping into shaped bodies such as tablets, strands and spheres and calcining the shaped bodies, catalyst supports with high mechanical stability and a produce a very suitable pore structure for the dehydrogenation of carbonyl compounds. The catalyst supports obtained are sufficiently stable to withstand several hundred oxidative regeneration cycles without mechanical damage and loss of activity.

The organosilicon compounds used as binders are generally liquid. As a result, the high-surface area zirconium dioxide is uniformly wetted with the organosilicon compound during mixing, as a result of which the zirconium dioxide particles are enclosed by the organosilicon compound and partially soaked. This results in a high level of bond strengthening between the zirconium dioxide particles and very good mechanical stability of the catalyst support moldings obtained. When the shaped catalyst support bodies are calcined, the organic residues of the organosilicon binder burn. This forms SiO ₂ , which is very finely distributed in the zirconium dioxide matrix. The combustion of the organic residues of the organosilicon binder creates additional pores. These pores are also very evenly distributed due to the even distribution of the organosilicon binder in the zirconia matrix. This increases the overall porosity of the catalyst carrier. The presence of SiO ₂ also stabilizes the zirconium dioxide against thermal sintering. This is more pronounced the more evenly the silicon dioxide is distributed.

Suitable organosilicon binders are monomeric, oligomeric or polymeric silanes, alkoxysilanes, aryloxyilanes, acyloxysilanes, oximinosilanes, halosilanes, aminoxysilanes, aminosilanes, amidosilanes, silazanes or silicones, as described, for example, in Ullmann's Encyclopedia of Industrial Chemistry, pages 21 A24 to 56 are described. These include in particular the monomeric compound of the general formulas (I) to (VI) below:

(HarjxSilU-x (I)

(Hal) _x Si (OR ¹ ) ₄ - _x (II) (Hal) _x Si (NR ¹ R ² ) ₄ .χ (HI)

R _x Si (NR ¹ R ² ) ₄ - _x (V)

(R ¹ O) _x Si (NR ¹ R ² ) ₄ -χ (VI)

wherein

Shark independently halogen (F, Cl, Br or I),

R independently of one another H or an optionally substituted alkyl, alkenyl, alkynyl, cycloalkyl, cycloalkenyl, arylalkyl or aryl radical, R, R independently of one another H or an optionally substituted alkyl, acyl, arylalkyl or aryl radical, and x mean 0 to 4.

R, R ¹ and R ² can denote H, an alkyl radical, preferably a Ci to C ₆ alkyl radical, which can be linear or branched. If R is an alkyl radical, R is in particular methyl, ethyl, n-propyl, isopropyl, n-butyl, sec-butyl or tert-butyl, especially methyl or ethyl. R, R ¹ and R ² can furthermore be an aryl radical, preferably phenyl, or an arylalkyl radical, preferably benzyl.

R can also mean an alkenyl radical, preferably a C ₂ -C ₆ alkenyl radical, in particular vinyl or allyl, or an alkynyl radical, preferably ethynyl.

R ¹ and R ² can also mean an acyl radical, preferably a C ₂ -C ₆ acyl radical, in particular an acetyl radical.

Examples of suitable organosilicon compounds of the general formula (I) are SiCl ₄ , MeSiCl ₃ , Me ₂ SiCl ₂ and Me ₃ SiCl.

Suitable organosilicon compounds of the general formula (IN) are, for example, Si (OEt) ₄ , MeSi (OEt) ₃ , Me ₂ Si (OEt) ₂ and Me ₃ SiOEt.

Suitable compounds of the general formula (V) are, for example, Me ₃ Si (ΝMeCOMe) and MeSi (NMeCOCH ₂ C ₆ H ₅ ).

A suitable compound of the general formula (VI) is, for example, (MeO) ₃ SiNMe ₂ . Examples of suitable oligomeric and polymeric organosilicon compounds are methyl silicones and ethyl silicones.

Very particularly preferred Wacker organosilicon binder are methyl silicones, for example, the Silres ^® brands from..

In a first step, zirconium dioxide powder is mixed with the organosilicon binder, optionally a pore former, optionally an acid, water and optionally further additives to form a kneadable mass. Preferably be

a) 50 to 98% by weight of zirconium dioxide powder, b) 2 to 50% by weight, particularly preferably 5 to 20% by weight of the organosilicon compound, c) 0 to 48% by weight, particularly preferably 0 to 10% by weight. % Pore-forming agents, and d) 0 to 48% by weight, particularly preferably 0 to 10% by weight, of further additives,

the sum of components a) to d) being 100% by weight,

mixed with the addition of water and an acid to a kneadable mass.

The zirconia powder has a high surface area, usually it is essentially monoclinic zirconia powder. Essentially monoclinic zirconium dioxide powder, which consists of 85 to 100% by weight, preferably 90 to 100% by weight, of monoclinic zirconium dioxide can be prepared as described in EP-A 0 716 883 by precipitation of zirconium salts with ammonia by: adding a zirconyl nitrate or zirconyl chloride solution to an aqueous ammonia solution, the pH falling from 14 to 6, the precipitate being washed out, dried and calcined. To this end, zirconium carbonate and hydrochloric acid are first prepared as concentrated as possible, usually 2 to 5 mol% zirconium chloride solution, or preferably zirconium carbonate and nitric acid are used to prepare a concentrated, usually 2 to 5 mol% zirconium nitrate solution. This solution is generally added to a submitted aqueous ammonia solution (approx. 15 mol% NH ₃ ) at temperatures of 20 to 60 ° C. while checking the pH, the addition being carried out at a pH of 6 to 8 is ended and the pH must not drop below 6. This is followed by a subsequent stirring time of generally 30 to 600 minutes. The precipitate is washed out, for example, on a filter press and essentially freed of ammonium salts, dried and calcined in air at a temperature of 300 to 600 ° C., preferably 400 to 500 ° C. and a pressure of 0.05 to 1 bar. Occasionally, the essentially monoclinic zirconia so produced still contains small amounts of the tetragonal or cubic modification. The proportion of the tetragonal or cubic modification can be reduced to the x-ray detection limit if drying is carried out under a water vapor partial pressure of 0.2 to 0.9 bar before the calcination. Drying takes about 16 hours at 120 ° C, for example.

Water is usually added to the zirconium dioxide powder and the organosilicon compound in order to obtain a kneadable mass.

An acid can also be added to the catalyst support composition. This causes the modeling clay to be peptized. Suitable acids are, for example, nitric acid and acetic acid, nitric acid is preferred.

The catalyst support mass usually contains a pore former. Suitable pore formers are, for example, polyalkylene oxides such as polyethylene oxide, carbohydrates such as cellulose and sugar, natural fibers, pulp or synthetic polymers such as polyvinyl alcohol.

The catalyst support molding composition may also contain other additives. Other additives are, for example, known compounds which influence the rheology.

Components a) to f) are mixed and homogenized in conventional mixing apparatus. Suitable mixing apparatuses are, for example, kneaders, pan mills and mix-mullers, which ensure thorough mixing and homogenization of the initially inhomogeneous kneadable mass. The catalyst support molding composition is then shaped to give shaped articles, for example by extrusion into strands or hollow supports.

The shaped catalyst support bodies are usually subsequently dried. The drying is carried out, for example, at a temperature of 90 to 120 ° C. over a period of 10 to 100 hours. The dried shaped catalyst support body is then calcined. The calcination is usually carried out at a temperature of 300 to 800 ° C, preferably from 400 to 600 ° C over a period of 0.5 to 6 hours. The calcination is preferably carried out in air and at atmospheric pressure.

The catalyst supports impregnated with the relevant metal salt solutions are usually calcined at a temperature of 350 to 650 ° C. over a period of 0.5 to 6 hours.

Regeneration of the dehydrogenation catalyst

In the dehydrogenation of carbonyl compounds according to the invention, small amounts of high-boiling, high molecular weight organic compounds or carbon are formed over time, which deposit on the catalyst surface and in the pores and deactivate the catalyst over time.

Used dehydrogenation catalysts are usually regenerated by purging with an inert gas, passing through an oxygen-containing gas mixture, purging with an inert gas and subsequent activation with hydrogen, working at atmospheric pressure. In the process according to US Pat. No. 5,087,792, the catalyst is regenerated by purging with inert gas, passing through an oxygen-containing gas mixture, purging with inert gas and then passing through an HCl / oxygen mixture to redisperse the active metal (palladium) on the support.

The regeneration process preferably comprises the following step (b): (b) Passing through an oxygen-containing gas mixture containing an inert gas at a pressure of 0.5 to 20 bar and a gas load of 1000 to 50,000 h ^"1 over a period of 0.25 to 24 h with a gradual or continuous increase in the oxygen concentration from an initial value of 0.01 to 1% by volume of O ₂ to a final value of 10 to 25% by volume of O ₂ .

Before step (b) is carried out, step (a) is usually carried out first:

(a) Purging with inert gas at a pressure of 0.5 to 2.0 bar and a gas load of 1000 to 50,000 h ^"1 .

If necessary, steps (c) and / or (d) are then carried out: (c) optionally passing an oxygen-containing gas mixture containing an inert gas at a pressure of 0.5 to 20 bar and a gas load of 10 to 500 h ^"1 over a period of 0.25 to 100 h, the oxygen concentration being 10 to 25 vol. % O ₂ ;

(d) optionally repeated rapid opposite pressure changes by a factor of 2 to 20 within the range of 0.5 to 20 bar.

After step (b), (c) or (d), step (e) is usually carried out:

(e) purging with an inert gas or water vapor.

Step (f) is preferably carried out last:

(f) Activation of the catalyst with hydrogen.

Usually at least one of steps (c) or (d) is carried out and the entire regeneration process is carried out at a temperature of 300 to 800 ° C. Steps (b) and optionally (c) are preferably carried out at a temperature above 500 ° C.

The dehydrogenation catalyst is preferably installed in a dehydrogenation reactor. However, it can also be regenerated in a separate regeneration reactor.

In step (a) the purging is preferably carried out with inert gas until the purging gas essentially contains no traces of the dehydrogenation product and hydrogen anymore, that is to say such traces can no longer be detected using the usual analytical methods, for example gas chromatography. In general, purging over a period of 0.1 to 24 h is required at a pressure of 0.5 to 2.0 bar and a gas load of 1000 to 50 000 h ^-1 . The pressure is preferably 1 to 1.5 bar, the gas load preferably 2000 to 20,000 h ^"1 . The duration of the rinsing step is preferably 0.1 to 6 hours. Nitrogen is generally used as the inert gas. The purge gas can also contain water vapor in amounts of, for example, 10 to 90% by volume.

In step (b), an oxygen-containing gas mixture is passed through the catalyst bed in order to burn off the surface coke deposits on the catalyst particles. Diluted air is preferably used as the oxygen-containing gas mixture, which in addition to an inert gas can also contain water vapor, for example in amounts of 10 to 90% by volume. The oxygen content is gradually increased, in generally from an initial concentration of 0.01 to 1% by volume, for example 0.1% by volume, to a final concentration of 10 to 25% by volume. If the oxygen-containing gas contains no water vapor and air is used, the final concentration is generally approx. 21% by volume. Oxygen. It is essential that work is carried out at a pressure significantly above the pressure that prevails during the dehydrogenation. The pressure is preferably 3 to 7 bar, for example 4 to 6 bar. The treatment time is preferably 0.5 to 12 h, for example 1 to 9 h. In general, a high gas load is used. This is preferably 2,000 to 20,000 h ^"1 .

In step (c), an oxygen-containing gas mixture which has a high oxygen content is passed through the catalyst bed. Air is preferably used for this. The oxygen-containing gas mixture can contain water vapor, for example in amounts of 10 to 90% by volume. In this step, the catalyst particles of coke deposited in the pores are burned off. This is carried out at low gas loads of generally 10 to 500 h ^"1 , preferably 20 to 100 h ^{" 1} . The pressure is not critical, it can be the same as the pressure in step (b) or lower. In general, it is 0.5 to 20 bar, preferably 1 to 5 bar.

In step (d) the pressure is repeatedly changed in the opposite direction, that is to say the pressure increases and the pressure decreases in short time intervals. In this way, CO formed in the pores can be effectively removed. Preferably, the pressure is changed 2 to 20 times by a factor of 2 to 5 within the range of 1 to 5 bar. For example, a total of 3 pressure increases and decreases from 1 to 5 and 5 to 1 bar are carried out. The total duration of all pressure increase and pressure reduction steps is preferably 0.1 to 1 h. In order for the pressure in the reactor to build up quickly, the gas load should not be chosen too low and is generally 100 to 50,000 h ^"1 , preferably 1000 to 20,000 h ^{" 1} .

Step (c) and step (d) can alternatively be carried out, usually at least one of these steps is carried out. Step (d) is carried out in particular when step (c) has been carried out over a short period of time, for example 0.25 to 5 hours. If step (c) is carried out over a longer period of time, for example 20 to 100 h, step (d) can be omitted.

In step (e), flushing is carried out with inert gas such as nitrogen or argon or water vapor, preferably over a period of 1 min to 1 h. This is followed in step (f) by the known activation of the catalyst with hydrogen. You can use pure hydrogen or with a hydrogen-containing gas, which can contain an inert gas and / or water vapor, for example in amounts of 10 to 90% by volume. The activation is preferably carried out at normal pressure over a period of 10 minutes to 2 hours.

In all steps (a) to (f), the temperature is generally 300 to 800 ° C, preferably 400 to 700 ° C. Steps (b) and (c) are preferably carried out at a temperature above 500 ° C.

dehydration

The gas phase dehydrogenation of the carbonyl compound takes place as a so-called non-oxidative dehydrogenation. The corresponding saturated carbonyl compound is at least partially dehydrogenated in a dehydrogenation reactor over the dehydrogenation-active catalyst to give the alpha, beta-unsaturated carbonyl compound, it also being possible for polyunsaturated products to be formed if appropriate. In addition, hydrogen and, in small quantities, low-molecular-weight by-products such as methane, ethane, ethene, propane and propene are obtained. Depending on how the dehydrogenation is carried out, carbon oxides (CO, CO ₂ ), water and nitrogen can also be present in the product gas mixture. In addition, unreacted starting material is generally also present in the product gas mixture.

Cyclic or acyclic aldehydes or ketones can be dehydrogenated as cyclic or acyclic carbonyl compounds. Examples of acyclic aldehydes and ketones which can be dehydrogenated to the corresponding alpha, beta-unsaturated compounds by the process according to the invention are propionaldehyde, butyraldehyde, valeraldehyde, iso-valeraldehyde, butanone, 2-pentanone and 2-hexanone. The dehydrogenation follows the following general reaction scheme:

where R ¹ H or methyl, R ² H, methyl or ethyl and R ³ and R ⁴ independently of one another H, Ci-GrAlkyl (methyl, ethyl, 1-propyl, 2-propyl, n-butyl, iso-butyl, sec- Butyl, tert-butyl) or optionally substituted phenyl or pyridyl.

As cyclic ketones, for example, cyclopentanone, cyclohexanone and cycloheptanone can be dehydrated by the process according to the invention. The dehydrogenation of the cyclic ketones follows the general reaction scheme:

where Y is -CH ₂ -, -CH ₂ CH ₂ - or -CH ₂ CH ₂ CH ₂ -.

The non-oxidative catalytic dehydrogenation can be carried out with or without oxygen-containing gas as an additional feed gas stream.

A feature of the non-oxidative mode of operation compared to an oxidative mode of operation is the presence of hydrogen in the discharge gas. In the oxidative dehydrogenation, free hydrogen is not produced in substantial amounts.

The non-oxidative catalytic dehydrogenation can in principle be carried out in all reactor types and procedures known from the prior art. A comparatively comprehensive description of dehydrogenation processes suitable erfmdungsgemäß also contains "Catalytica® ^® Studies Division, Oxidative Dehydrogenation and Alternative Dehydrogenation Processes" (Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, California, 94043-5272, USA).

A suitable reactor form is the fixed bed tube or tube bundle reactor. These contain the catalyst (dehydrogenation catalyst and, when working with oxygen as an additional feed gas stream, possibly a special oxidation catalyst) as a fixed bed in a reaction tube or in a bundle of reaction tubes. The reaction tubes are usually heated indirectly in that the A gas, for example a hydrocarbon such as methane, is burned or a heat transfer medium (salt bath, rolling gas etc.) is used. The reaction tubes can also be heated electrically with heating jackets. Usual reaction tube inner diameters are about 10 to 15 cm. A typical dehydrogenation tube bundle reactor comprises approximately 10 to 10,000 reaction tubes, preferably approximately 10 to 200 reaction tubes. The temperature in the interior of the reaction tube is usually in the range from 300 to 1200 ° C., preferably in the range from 400 to 600 ° C. The working pressure is usually between 0.5 and 8 bar, often between 1 and 2 bar when using a low water vapor dilution (analogous to the Linde process for propane dehydrogenation), but also between 3 and 8 bar when using a high water vapor dilution (analogous to that So-called "steam active reforming process" (STAR process) for the dehydrogenation of propane or butane from Phillips Petroleum Co., see US 4,902,849, US 4,996,387 and US 5,389,342). Typical catalyst loads (GHSV) are 500 to 2000 h ^"1 on hydrocarbon used. The catalyst geometry can be spherical or cylindrical (hollow or full), for example.

The non-oxidative catalytic dehydrogenation can also, as in Chem. Eng. Be. 1992 b, 47 (9-11) 2313, heterogeneously catalyzed in a fluidized bed. Appropriately, two fluidized beds are operated side by side, one of which is usually in the state of regeneration. The working pressure is typically 1 to 2 bar, the dehydrogenation temperature usually 550 to 600 ° C. The heat required for the dehydrogenation is introduced into the reaction system by preheating the dehydrogenation catalyst to the reaction temperature. By adding an additional feed gas stream (co-feed) containing oxygen, the preheaters can be dispensed with and the required heat can be generated directly in the reactor system by burning hydrogen in the presence of oxygen. If necessary, a hydrogen-containing co-feed can also be added.

The non-oxidative catalytic dehydrogenation can be carried out with or without oxygen-containing gas as a co-feed in a tray reactor. This contains one or more successive catalyst beds. The number of catalyst beds can be 1 to 20, advantageously 1 to 6, preferably 1 to 4 and in particular 1 to 3. The reaction beds preferably flow radially or axially through the catalyst beds. Such a tray reactor is generally operated with a fixed catalyst bed. In the simplest case, the fixed catalyst beds are arranged axially in a shaft furnace reactor or in the annular gaps of concentrically arranged cylindrical gratings. A shaft furnace reactor corresponds to a horde. The implementation of the dehydrogenation in a single shaft furnace reactor corresponds to a preferred embodiment, it being possible to work with an oxygen-containing co-feed. In a further preferred embodiment, the dehydrogenation is carried out in a tray reactor with 3 catalyst beds. In a mode of operation without oxygen-containing gas as a co-feed, the reaction gas mixture in the tray reactor is subjected to intermediate heating on its way from one catalyst bed to the next catalyst bed, for example by passing it over heat-exchanger surfaces heated with hot gases or by passing it through pipes heated with hot fuel gases.

In one embodiment of the process according to the invention, the non-oxidative catalytic dehydrogenation is carried out autothermally. For this purpose, oxygen is additionally added to the reaction gas mixture in at least one reaction zone and the hydrogen and / or hydrocarbon contained in the reaction gas mixture is at least partially burned, as a result of which at least some of the heat of dehydrogenation required is generated directly in the reaction gas mixture in the at least one reaction zone.

In general, the amount of the oxygen-containing gas added to the reaction gas mixture is selected such that the amount of heat required for the dehydrogenation is generated by the combustion of hydrogen present in the reaction gas mixture and optionally of the carbonyl compounds present in the reaction gas mixture and / or of carbon present in the form of coke. In general, the total amount of oxygen supplied, based on the total amount of carbonyl compounds, is 0.001 to 0.5 mol / mol, preferably 0.005 to 0.2 mol / mol, particularly preferably 0.05 to 0.2 mol / mol. Oxygen can be used either as pure oxygen or as an oxygen-containing gas in a mixture with inert gases, for example in the form of air. The inert gases and the resulting combustion gases generally have an additional dilution effect and thus promote heterogeneously catalyzed dehydrogenation. The hydrogen burned to generate heat is the hydrogen formed in the catalytic dehydrogenation and, if appropriate, hydrogen additionally added to the reaction gas mixture as a hydrogen-containing gas. Sufficient hydrogen should preferably be present so that the molar ratio H * 2 O ₂ in the reaction gas mixture is 1 to 10, preferably 2 to 5 mol / mol immediately after the oxygen has been fed in. In the case of multi-stage reactors, this applies to every intermediate feed of oxygen-containing and possibly hydrogen-containing gas. The hydrogen is burned catalytically. The dehydrogenation catalyst used generally also catalyzes the combustion of the carbonyl compound and of hydrogen with oxygen, so that in principle no special oxidation catalyst different from this is required. In one embodiment, the process is carried out in the presence of one or more oxidation catalysts which selectively catalyze the combustion of hydrogen to oxygen in the presence of the carbonyl compound. The combustion of the carbonyl compound with oxygen to CO, CO ₂ and water takes place only to a minor extent. The dehydrogenation catalyst and the oxidation catalyst are preferably present in different reaction zones.

If the reaction is carried out in several stages, the oxidation catalyst can be present in only one, in several or in all reaction zones.

The catalyst, which selectively catalyzes the oxidation of hydrogen, is preferably arranged at the points where there are higher oxygen partial pressures than at other points in the reactor, in particular in the vicinity of the feed point for the oxygen-containing gas. Oxygen-containing gas and / or hydrogen-containing gas can be fed in at one or more points in the reactor.

In one embodiment of the method according to the invention, there is an intermediate feed of oxygen-containing gas and of hydrogen-containing gas upstream of each tray of a tray reactor. In a further embodiment of the method according to the invention, oxygen-containing gas and hydrogen-containing gas are fed in before each horde except the first horde. In one embodiment, a layer of a special oxidation catalyst is present behind each feed point, followed by a layer of the dehydrogenation catalyst. In a further embodiment, no special oxidation catalyst is present. The dehydrogenation temperature is generally 400 to 1100 ° C, the pressure in the last catalyst bed of the tray reactor generally 0.2 to 5 bar, preferably 1 to 3 bar. The load (GHSV) is generally 500 to 2000 h ^"1 , in the high-load mode also up to 100,000 h ^{" 1} , preferably 4000 to 16000 h ^"1 .

A preferred catalyst that selectively catalyzes the combustion of hydrogen contains oxides and / or phosphates selected from the group consisting of the oxides and / or phosphates of germanium, tin, lead, arsenic, antimony or bismuth. Another preferred catalyst which catalyzes the combustion of hydrogen contains a noble metal of the VIII. And / or I. subgroup. Processing of the dehydrogenation discharge

The dehydrogenation output can be worked up continuously or batchwise.

The reaction discharge from the dehydrogenation essentially consists of the alpha, beta-unsaturated carbonyl compound, unreacted starting material, water, hydrogen, CO, CO ₂ and low-boiling hydrocarbons such as methane, ethene, ethane, propene and propane. Low-boiling constituents of the reactor discharge can be separated off in a condenser, a mixture of water, product and starting material being obtained as the condensate. Because of the high polarity of the educt and product carbonyl compounds, there is often no or only poor phase separation into organic and aqueous phases. In this case, the work-up can comprise an extraction of the aqueous condensate with an organic extracting agent. An extracting agent is preferably used which has either a significantly higher or a significantly lower boiling point than the starting material and product. The extractant is then recovered in a distillation. The remaining educt / product mixture is then separated in a further distillation. The recovered starting material is preferably returned to the dehydrogenation. If there are azeotropes between water and the starting material and / or product, the extracting agent is preferably selected such that it also entrains the water remaining in the organic phase when it is distilled off. Alternatively, instead of extraction, water can also be separated off by azeotropic distillation with a suitable entrainer, such as, for example, cyclohexane.

Working up can be carried out as shown in FIG. I for the dehydrogenation of cyclopentanone to give 2-cyclopenten-1-one as an example. Cyclopentanone (1) and water / steam mixture (2) are fed to the evaporator (3) and evaporated together. Hydrogen (4) and optionally oxygen or an oxygen-containing gas (5) are fed into the gaseous cyclopentanone / water vapor mixture and the gas mixture is fed to the dehydrogenation reactor (6). The reactor discharge (7) essentially contains the dehydrogenation product 2-cyclopenten-l-one, unreacted cyclopentanone, water vapor, hydrogen, CO, CO ₂ and low-boiling hydrocarbons such as methane. Low boilers (9) such as hydrogen, CO, CO and methane are separated off in the condenser (8). The condensate (10) consists essentially of water, 2-cyclopenten-l-one and cyclopentanone. In a subsequent extraction stage (11), the organic constituents of the condensate (10) are extracted with an organic extractant (15). Suitable organic extractants are, for example, ethyl acetate, methyl tert.- buytlether and dichloromethane. The aqueous phase leaving the extraction stage (11) can at least partially be returned to the dehydration as a partial stream (12a). The organic phase (13) leaving the extraction stage, which essentially consists of 2-cyclopenten-1-one, cyclopentanone and organic extractant, is fed to the distillation column (14), in which the organic extractant, for example at a pressure of 150 mbar, withdrawn overhead and thus recovered. In order to avoid the accumulation of by-products, a partial stream (15a) ("purge stream") of the recovered extractant is separated from the main stream (15). The bottom draw stream (16. Essentially consisting of cyclopentanone and 2-cyclopenten-1-one) ) is separated in a further distillation column (17), preferably under reduced pressure, for example at 70 mbar, into cyclopentanone 18 as the top product and 2-cyclopenten-l-one (19) as the bottom product, and the cyclopentanone 18 is returned to the dehydrogenation Crude product (19) is further purified by distillation in a downstream purification nozzle column (20), the pure product being obtained as top draw stream 21. The purification nozzle column (20) is preferably operated at reduced pressure, for example of 30 mbar Bottom liquefier (22), for example propylene carbonate or fatty alcohol ethoxylates, is added FLOWING as a bottom draw stream (23) ( "purge stream") removed.

The invention is illustrated by the examples below.

Examples

example 1

Preparation of the dehydrogenation catalyst

1000 g of a split ZrO ₂ ^β SiO ₂ mixed oxide from Norton (sieve fraction 1.6 to 2 mm) were mixed with a solution of 11.992 g SnCl ₂ ⁹ 2H ₂ O and 7.888 g H ₂ PtCl ₆ * 6H ₂ O in 5950 ml Poured ethanol.

The supernatant ethanol was removed on a rotary evaporator. The mixture was then dried at 100 ° C. for 15 h and calcined at 560 ° C. for 3 h. The catalyst obtained was then poured with a solution of 7.68 g of CsNO ₃ , 13.54 g of KNO ₃ and 98.329 g of La (NO ₃ ) ₃ * 6H ₂ O in 23 ml of H ₂ O. The supernatant water was drawn off on a rotary evaporator. The mixture was then dried at 100 ° C. for 15 h and calcined at 560 ° C. for 3 h. The catalyst had a BET surface area of 85 m Ig. Mercury porosimetry measurements showed a pore volume of 0.29 ml / g.

Example 2 Dehydrogenation of cyclopentanone to 2-cyclopentenone

Cyclopentanone and water are evaporated together via an evaporator and dehydrated in a continuous process in a tubular reactor at 500 ° C. over the catalyst from Example 1. The mass ratio of the water vapor: cyclopentanone feeds is 1: 1. The catalyst is loaded with an LHSV of 1.25 / h. The LHSV is defined as the cyclopentanone flow (defined as the liquid volume flow under standard conditions) per volume of the catalyst bed.

The reactor discharge is liquefied on a cooler at 0 ° C and separated from non-condensed gaseous components. The liquefied reaction product, which contains about 50% water, is extracted with ethyl acetate. After separation of the two phases, the organic phase, which consists essentially of the extractant, cyclopentanone and cyclopentenone, is subjected to a fractional distillation.

First, the extractant (ethyl acetate) is distilled off at 150 mbar and 40 ° C. under reduced pressure. The high boiler fraction drawn off at the bottom is mixed with bottom liquefier. In the subsequent distillation stages, cyclopentanone at 70 mbar and 60 ° C. and finally 2-cyclopenten-l-one at 30 mbar and 60 ° C. are separated off from this at the top under reduced pressure.

The cyclopentenone obtained has a purity of> 99.5%. The recovered cyclopentanone is used again for dehydrogenation. The extractant is returned to the extraction.

In an 8-hour test run, an average conversion of 18%, falling from approx. 24% after 1 h to approx. 8% after 8 h, was achieved with an average selectivity of the formation of 2-cyclopenten-l-one of 85% , In addition to organic deposits on the catalyst, gaseous compounds such as methane, ethylene, propylene and carbon oxides were formed practically exclusively as a by-product.

The catalyst is regenerated at regular intervals. Example 3

The example was carried out analogously to Example 2, but with an LHSV of 0.4 h ^"1 and a water vapor content of 50% by weight, based on cyclopentanone. The discharges were collected and analyzed after 8 hours. A conversion of 24 , 8% found.

Claims

claims

1. A process for the preparation of alpha, beta-unsaturated acyclic or cyclic carbonyl compounds by dehydrogenation of the corresponding saturated carbonyl compounds in the gas phase on a heterogeneous dehydrogenation catalyst, characterized in that the dehydrogenation catalyst contains platinum and / or palladium and tin on an oxidic support.

2. The method according to claim 1, characterized in that the oxidic carrier is selected from the group consisting of zirconium dioxide, aluminum oxide, silicon dioxide, titanium dioxide, magnesium oxide, lanthanum oxide and cerium oxide.

3. The method according to claim 1 or 2, characterized in that the dehydrogenation catalyst contains zirconium dioxide and / or silicon dioxide.

4. The method according to any one of claims 1 to 3, characterized in that the dehydrogenation catalyst additionally at least one element of the I. or II. Main group and at least one element of III. Subgroup, including the lanthanides and actinides, contains.

5. The method according to any one of claims 1 to 4, characterized in that the dehydrogenation catalyst contains cesium and / or potassium.

6. The method according to any one of claims 1 to 5, characterized in that the dehydrogenation catalyst contains lanthanum and / or cerium.

7. The method according to any one of claims 1 to 6, characterized in that the dehydrogenation is carried out in the presence of molecular oxygen under autothermal conditions.

8. The method according to any one of claims 1 to 7, characterized in that a regenerated dehydrogenation catalyst is used, the regeneration comprising at least the following step: passing an oxygen-containing gas mixture containing an inert gas through the catalyst bed at a pressure of 2 to 20 bar and one Gas pollution of 1000 - 2 - to 50,000 h ^"1 over a period of 0.25 to 24 h with gradual or continuous increase in the oxygen concentration from an initial value of 0.01 to 1 vol .-% O ₂ to a final value of 10 to 25 vol .-% O ₂ .

9. The method according to claim 8, characterized in that the regeneration of the dehydrogenation catalyst comprises steps (a), (b) and (e) and optionally (c), (d) and (f):

(a) flushing with inert gas at a pressure of 0.5 to 2.0 bar and a gas load of 1000 to 50,000 h ^"1 ; (b) passing an oxygen-containing gas mixture containing an inert gas through the catalyst bed at a pressure of 2 to 20 bar and a gas load of 1000 to 50 000 h ^"1 over a period of 0.25 to 24 h with gradual or continuous increase in the oxygen concentration from an initial value of 0.01 to 1% by volume of O ₂ to a final value from 10 to 25 vol% O ₂ ; (c) optionally passing an oxygen-containing gas mixture containing an inert gas at a pressure of 0.5 to 20 bar and a gas load of 10 to 500 h ^"1 over a period of 0.25 to 100 h, the oxygen concentration being 10 to 25 vol. -% O ₂ ; (d) optionally repeated rapid opposite pressure change by a factor of 2 to 20 within the range from 0.5 to 20 bar; (e) flushing with an inert gas; (f) activation of the catalyst with hydrogen; wherein at least one of steps (c) or (d) is carried out and the entire regeneration process is carried out at a temperature of 300 to 800 ° C.

10. The method according to claim 8 or 9, characterized in that steps (b) and optionally (c) are carried out at a temperature of> 500 ° C.

11. The method according to any one of claims 1 to 10, characterized in that the cyclic or acyclic carbonyl compounds are cyclic or acyclic aldehydes or ketones. - 3 -

12. The method according to claim 11, characterized in that an acyclic aldehyde or ketone selected from the group consisting of propionaldehyde, butyraldehyde, valeraldehyde, iso-valeraldehyde, butanone, 2-pentanone and 2-hexanone is dehydrogenated.

13. The method according to claim 11, characterized in that a cyclic ketone selected from the group consisting of cyclopentanone, cyclohexanone and cycloheptanone is dehydrated.

14. Use of the regenerated dehydrogenation catalyst as defined in one of claims 8 to 10 in the process according to claim 1.