MXPA00011539A - Process for the hydroformylation of olefins - Google Patents

Abstract

Aldehydes are prepared in a hydroformylation apparatus by the multiphase hydroformylation reaction of one or more olefins with hydrogen and carbon monoxide, where the continuous phase contains a solvent mixture and the hydroformylation catalyst is present in the continuous phase, at least one olefin is present in the dispersed phase, and the loading factor of the tube reactor is greater than or equal to 0.8.

Description

Process for the hydroformylation of olefins The present invention relates to a process for the preparation of aldehydes of 3 to 26 carbon atoms by the reaction of olefins of 2 to 25 carbon atoms with hydrogen and carbon monoxide in the presence of a catalyst in a tubular reactor. Aldehydes are used for the synthesis of many organic compounds. Its direct sequential products are alcohols and carboxylic acids that in turn are technically exploited. By means of the aldol condensation of these aldehydes and the subsequent hydrogenation of the condensate, alcohols with twice as many carbon atoms are created as starting aldehydes. Among other purposes, the alcohols prepared from the aldehydes by hydrogenation are used as a solvent and as a preliminary stage for the preparation of plasticizers and detergents. The preparation of aldehydes and alcohols is known (hydroformylation, oxo reaction) by the reaction of olefins with carbon monoxide and hydrogen. The reaction is catalyzed by hydrid metalic carbonyls, preferably by those of metals of the eighth group of the periodic system of the elements. Recently, along with cobalt which is widely used in the art as a catalyst metal, the importance of rhodium has increased. Compared with cobalt, rhodium allows the realization of the reaction at low pressures. The hydrogenation of the olefins to give saturated hydrocarbons is carried out using rhodium catalysts in a range considerably less than the use of cobalt catalysts. In hydroformylation processes carried out in the art, the rhodium catalyst is formed during the process from a catalyst precursor, synthesis gas and, optionally, other ligands. By using modified catalysts, the modified ligands may be present in excess in the reaction mixture. In particular, tertiary phosphines or phosphite ligands are suitable. Its use makes it possible for the reaction pressure to be lowered to values considerably lower than 300 bar. In this process, however, the separation of the reaction products and the recovery of the catalysts homogeneously dissolved in the reaction product presents problems. In general, the reaction product is distilled from the reaction mixture. In practice, the above is only possible in the hydroformylation of lower olefins of up to 5 carbon atoms in the molecule, due to the thermal sensitivity of the catalyst or of the products formed.

On an industrial scale, aldehydes of 4 and 5 carbon atoms are prepared, for example, according to DE 32 34 701 or DE 27 15 685. In the process according to DE 27 15 685, the catalyst is dissolved in a organic phase consisting of high boiling product and substance (generated from the product). In this mixture, the olefin and synthesis gas are introduced. With the synthesis gas, the product is extracted from the reactor, that is, it is decanted as a liquid. Because the catalyst slowly loses its effectiveness, a part must be continually removed together with the high-boiling substance and replaced with an equivalent amount of fresh catalyst. The recovery of rhodium from the extraction flow is indispensable, since its price is high. The processing process is complex and, therefore, complicates the procedure. In accordance with DE 32 34 701, this disadvantage is overcome, for example, by dissolving the catalyst in water. The water solubility of the rhodium catalyst used is obtained by trisulfonated triarylphosphoria as a ligand. Olefin and synthesis gas are introduced into the aqueous catalytic phase. The product obtained by the reaction forms a second aqueous phase. The aqueous phases are separated out of the reactor and the separated catalytic phase is conducted back to the reactor. This last process, which includes the advantageous separation of the catalyst, has a space-time efficiency lower than the procedures, wherein an aqueous organic phase with a dissolved catalyst is presented. The reason for the above is the varied solubility of olefins. While the olefins dissolve well in an organic solution or even in themselves form the aqueous organic phase in which the dissolved catalyst is found, the olefins practically do not dissolve in aqueous solutions. The solubility of the olefins in an aqueous solution which is itself low decreases in proportion to the increase in the molecular mass of the olefins. Consequently, higher aldehydes can not be prepared economically according to this procedure. By the addition of an organic solvent soluble in the aqueous catalytic phase, the reaction rate of the hydroformylation can be increased. The use of alcohols, such as methanol, ethanol or isopropanol as co-solvent increases the reaction rate, however, it has the disadvantage that rhodium is transformed into the product phase (B. Cornils, A. Herrmann, Aqueous-Phase Organometallic Catalysis, iley-VCH, page 316-317) and in this way is removed from the catalytic circuit. An increase in the reaction rate can, for example, be achieved by the addition of ethylene glycol. However, the selectivity of aldehyde formation decreases by this measure, since acetal derivatives are generated from aldehyde and ethylene glycol (Nair VSR, BM Bhanage, RM Deshpande, RV Chaudhari, Recent Advances in Basic and Applied Aspects of Industrial Catalysis, Studies in Surface Science and Cayalysis, Vol. 113, 529-539, 1998 Elevier Science BV). Patent EP 0 157 316 describes the increase in the reaction rate in the hydroformylation of 1-hexene with the addition of solubilizers such as carboxylic acid salts, glycollene alkylpolyethylene or quaternary onium compounds. Here, productivity was increased by factor 4 depending on the solubilizer. The increase in the rate of the reaction by the addition of polyglycollene (for example, PEG 400) and polyglycol ethers is known. Thus, patent DE 197 00 805 Cl describes the hydroformylation of propene, 1-butene and 1-pentene and patent DE 197 00 804 Cl describes the hydroformylation of higher olefins such as 1-hexene, 4-vinylcyclohexene, 1- octene, 1-decene or 1-dodecene. These processes have in common that by means of the use of solubilizers the reaction rate is increased, however, the separation of the aqueous catalytic phase and the organic phase of the product is also difficult. This means losses of catalyst that undesirably transforms from the aqueous phase into the organic phase, as well as loss of valuable products that become soluble in the aqueous phase. If the amount of solubilizer is decreased in order to minimize these losses, the speed of the reaction is reduced at the same time. DE 199 25 384 describes that it can improve the space-time yield of aldehydes in the hydroformylation of olefins in a polyphase reaction, where a continuous catalytic phase and an additional aqueous phase are present, if the reaction is not carried out in a agitator reactor but in a flow reactor with a load factor B >; 0.8. This process for the hydroformylation of olefins by a polyphase reaction has high load factors on the tubular reactor, that is, an extremely high mixing of the phases. To the catalytic phase, reagents or phase transfer surfactants, surfactants or amphiphiles can be added as an additive, while water is used as the preferred solvent for the catalyst.

Therefore, the present invention is based on the objective of developing a process for the hydroformylation of olefins which has a high space-time yield, as well as selectivities. Surprisingly it was found that the hydroformylation of olefins can be carried out in the form of a polyphase reaction with high yields and a lower formation of side products if the catalytic phase consists of a solvent mixture. Therefore, a process for the hydroformylation of one or more olefins of 2 to 25 carbon atoms by means of a polyphase reaction in a tubular reactor is provided, wherein a) the catalyst is contained in the continuous phase, b ) the continuous phase contains a solvent mixture, c) the dispersed phase contains at least one olefin and d) the load factor of the tubular reactor is equal to or greater than 0.8. In accordance with the present invention, the hydroformylation is carried out in a tubular reactor, that is, in a flow tube. The catalytic phase and the dispersed phase containing at least one olefin are pumped into the tubular reactor. After the reaction, the reaction mixture is separated into a product phase and a catalytic phase and the catalytic phase is conducted back to the tubular reactor. The product phase is removed from the circuit and can, for example, be processed by distillation to obtain the aldehydes. In addition, the use of the aldehydes prepared in this way is the object of the present invention. Thus, the aldehydes prepared according to the process of the present invention can be used for the preparation of alcohols by hydrogenation, in aldol condensations or for the preparation of carboxylic acids by oxidation. The catalytic solution used in the process according to the present invention contains a solvent mixture and a catalyst. As a component of the solvent mixture, protic-polar substances such as, for example, water, 1,2-propylene glycol, 1,3-propylene glycol, butanediols or glycerins can be used. A preferred solvent component is water. As an additional solvent component for forming the solvent mixture, polar organic substances can be used, in particular those containing at least two oxygen atoms. These are, for example, compounds of the classes of substances of the diols, triols, polyols and their partial or complete ethers. Examples are some compounds or groups of compounds: ethylene glycol, ethylene glycol monoether, ethylene glycol diether, ethylene glycol ethoxylates, ethylene glycol ethoxylate ether, ethylene glycol propoxylates, ethylene glycol propoxylate mono- and di-ether, Propylene glycol propoxylates, their mono- and di-ethers, polyols prepared by the hydrogenation of carbohydrates (for example, monosaccharides, disaccharides, hydrogenated oligosaccharides) and their partial and complete ethers. Therefore, in the process, according to the present invention, a solvent mixture consisting of water and an organic solvent miscible with water containing at least two carbon atoms can be employed as the continuous phase. The mass ratio of the solvent within the solvent mixture can vary within a wide range, as long as the following conditions are met: The resulting mixture should form a homogeneous phase. The solubility for the catalyst in this homogeneous solution (phase) should be sufficient for the desired catalyst concentrations. In addition, the solution should not become so viscous that difficulties in the reaction and / or during the subsequent separation of the phases occur.

Preferably, solvent mixtures having a dielectric constant of 50 to 78 at a temperature of 20 ° C are used. Examples for these solvent mixtures are water / ethylene glycol mixtures according to the following table: As hydroformylation catalysts, compounds of metals of the eighth secondary group of the periodic system of the elements can be used, such as Fe, Co, Ni, Ru, Rh, Pd, Os, Ir or Pt, preferably in the form of complex compounds. Advantageously, these metal compounds should be soluble in the catalytic phase, but not in the product phase. The use of aqueous catalyst solutions requires water-soluble metal compounds. Preferred catalysts are rhodium, that is, water-soluble rhodium compounds. Suitable rhodium salts are, for example, rhodium-III-sulfate, rhodium-III-nitrate, rhodium-III-carboxylate such as rhodium acetate, rhodiumpropionate, rhodiumbutyrate or rhodium-2-ethylhexanoate. The type of ligands for the metal complexes used as a catalyst depends on the metal and the solvent mixture used. These complexes should also not lose their catalytic effect during continuous use. A condition for the above is that the ligands are not transformed, for example, due to the reaction with the solvent. As ligands for the catalytic metals mentioned above, triarylphosphines can be used. Suitable phosphines have one or two phosphorus atoms which possess for each phosphorus atom three aryl residues, respectively, while the aryl residues may be the same or different and represent a phenyl, naphthyl, biphenyl, phenylnaphthyl or binaphthyl residue. The aryl moieties can be linked to the phosphorus atom directly or by means of a group - (CH2) x, where x represents an integer from 1 to 4, preferably from 1 to 2 and, of particular preference 1. For water-soluble catalyst systems, a ligand should contain one to three residues - (S03) M, where M can be the same or different and represents H, an alkali metal ion such as Na or K, a quaternary ammonium ion , medium (in the form of a calculation) alkaline earth metal ion such as Ca or Mag or a zinc ion. The residues -So3M are, in most cases, substituents in the aryl radicals and give the triarylphosphines the necessary water solubility. A preferred sulfonated triarylphosphine with the phosphorus atom is trisodium-tri- (m-sulfophenyl) phosphine. In the place of the sulfonate units (S03M), the phosphines used may also be substituted by other polar groups such as, for example, carboxylate units. In hydroformylation, the solvent mixture can be used directly, that is, without the catalyst, or the catalyst can be preformed in the solvent mixture and the mixture can be used with the preformed catalyst. If the solvent mixture contains water, the catalytic solution can be prepared in a similarly simple manner, by dissolving in water a water-soluble metal salt and / or the water-soluble ligands, leading to the formation of complexes, and then adding the solvent or the solvents for the formation of the solvent mixture. The concentration of the metal salt used in the process according to the present invention can be adjusted within a wide range, while the speed of the reaction also depends on the concentration of the metal salt. In general, higher reaction rates are achieved with higher concentrations of the metal salt. On the other hand, high concentrations of metallic salt also include higher costs. Therefore, depending on the reactivity of the starting material and the other reaction conditions, an optimum which is easily determined by orienting tests can be selected. In general, if rhodium is used as the active catalyst, the content of rhodium in the catalytic phase is 20 ppm up to 2000 ppm, preferably 100 up to 1000 ppm. The molar ratio between the metal and the ligands can vary in order to achieve the optimum for each of the reactions. This metal / ligand ratio is between 1/5 and 1/200, preferably between 1/10 and 1/60. The pH value of the catalytic solution can be optimized for the hydroformylation of each olefin with respect to the selectivity of the aldehyde formation. It is between 2 and 8, preferably between 3.0 and 5.5. The adjustment of the pH value can be carried out during the process and, for example, by the addition of caustic soda or sulfuric acid. In the process according to the present invention, olefinic compounds having 2 to 25 carbon atoms, preferably 3 to 12 carbon atoms, can be used as the educts. The olefinic compounds may contain one or more carbon-carbon double bonds which may be arranged in the terminal position or in the interior position, respectively. Preference is given to olefinic compounds with carbon-carbon double bond in terminal position. Here an olefin with a uniform structure can be used. The mixture may consist of isomeric olefins with the same number of carbon atoms or in olefins with a different number of carbon atoms or in a mixture containing both isomeric olefins and olefins with a different number of carbon atoms. In addition, under certain reaction conditions, olefins or olefin mixtures may contain inert substances such as hydrocarbons. Preferably, the olefins form, with the inert substances, the dispersed phase. In the process, according to the present invention, olefins of the most different origins can be used. As examples, olefins could be mentioned from cracking, dehydrogenation or from the Fischer-Tropsch synthesis. Also olefins or mixtures of olefins generated by dimerization, oligomerization, codimerization, cooligomerization or metathesis of olefins are suitable educts.

(Under normal conditions), the defines used can be presented in gaseous, liquid or solid form. The solid olefins can be used in the form of solutions. Solvents are inert liquids that are not soluble or poorly soluble in the catalytic phase. Particular preference is given to solvents which have a boiling point above the boiling point of the products to be prepared, since this facilitates separation by distillation and reflux. In the process according to the present invention preference is given to the use of -olefinic compounds. Examples for suitable α-olefinic compounds are 1-alkenes, alkyl alkenoates, alkene alkeneites, alkenyl alkyl ether and alkenols such as, for example, propene, butene, pentene, butadiene, pentadiene, 1-hexene, 1-heptene, 1- octene, 1-nonene, 1-decene, 1-undecene, 1-dodecene, 1-hexadecene, 2-ethyl-1-hexene, 1,4-hexadiene, 1,7-octadiene, 3-cyclohexyl-1-butene, styrene, 4-vinylcyclohexene, allylacetate, vinylformate, vinylacetate, vinylpropionate, allylmethyl ether, vinylmethyl ether, vinylethyl ether, allyl alcohol, 3-phenyl-1-propene, hex-l-en-4-ene, oct-l-en- 4-ene, 3-butenylacetate, allylpropionate, allylbutyrate, n-propyl-7-octenoate, 7-octene acid, 5-hexenamide, l-methoxy-2,7-octadiene and 3-methoxy-l, 7-octadiene.

In particular, propene, 1-butene or technically available olefin mixtures containing essentially 1-butene, 2-butene and i-butene, and / or 1-pentene are suitable. The products of the hydroformylation of olefins are aldehydes having one more carbon atom and, if appropriate, the corresponding alcohols which were generated by hydrogenation during the process according to the present invention. However, the aldehydes prepared according to the process, according to the present invention, can also be hydrogenated to give the corresponding saturated alcohols which can be used as solvents, as well as, for the preparation of detergents or plasticizers. In the process according to the present invention, mixtures of hydrogen and carbon monoxide (synthesis gas) or other mixtures of hydrogen, carbon monoxide and, under certain reaction conditions, inert substances are used as the hydroformylation agent. Preference is given to the use of synthesis gas with 50% by volume of H2 and 50% by volume of CO. When using liquid olefins or solid olefins in solution, it is advantageous if the hydroformylation agent is used in excess, so that the most complete reaction possible is achieved. This makes the processing more economical. When using gaseous olefins it may be advantageous if the hydroformylation reagent is used in excess, since the excess gaseous olefin is separated from the liquid phase of the product and can be introduced back into the process. The molar ratio of the olefin with the hydrogen or of the olefin with the carbon monoxide may be greater than, less than or equal to 1, respectively. When employing a gaseous olefin, the process, according to the present invention, is first a biphasic reaction, a liquid phase of product being formed during the reaction and, therefore, a three-phase system being created. When using a liquid olefin, an at least three-phase system is presented from the beginning. The tubular reactor used in the process according to the present invention may contain filler bodies or additional elements. The filler bodies, according to the present invention, are for example: Raschig rings, chair-shaped filler bodies, impact rings, saucers, wire mesh rings, wire lattices. Examples for additional elements are filter plates, flow cups, current switches, column plates, perforated plates or any other mixing device. As additional elements, according to the present invention, several narrow and parallel tubes can also be considered., resulting in a multi-tube reactor. Particular preference is given to mixer packs or dropper packs. In the method, according to the present invention, the conservation or the excess of the linear current density has a particular importance. In the ascending operation of the reactor (direction of the current from bottom to top) the flood point should be exceeded. That is to say, the reactor is operated in the upper margin at the point where bubble columns are generally operated. In the descending operation (direction of the current from top to bottom) the linear density of current must be adjusted in such a way that the reactor is completely flooded. Work is done in the upper margin at this point where there is still a trickle bed. To determine precisely the minimum reactor load that must be retained, the load factor B of the tubular reactor is calculated as a loss of pressure without dimension with B = PD / PS where PD [Pa / m] means a linear pressure loss at the length of the reactor under operating conditions and PS [Pa / m] means an operand with the dimension of a linear pressure, defined as the mass flow ratio M [kg / s] of all the components in the reactor with the volumetric flow rate V [ m3 / s] of all the components under operating conditions, multiplied by g = 9.81 m / s2, ie PS == (M / V) * g. Illustratively, PS would be the static pressure per meter in a multiphase mixture in a vertical tube if all phases flowed at the same velocity. PS is only an operand that results from the mass currents that are introduced into the reactor and that can be indicated independently of the direction of the reactor current, the speed of the current of all phases or the state of flooding of the reactor . The PD loss [Pa / m] is used as an operand to determine the process conditions and can be calculated according to the common methods for single-phase or multiphase currents. Common procedures for calculating the pressure loss PD in pipes, additional elements or layers of filler bodies, etc. can be investigated in the VDI-armament, 7th expanded edition, editorial VDI-Verlag GmbH, Düsseldorf, 1994, paragraphs Lal to Lgb7, as well as standard work Heinz Bauer Standard, Basic principles of single-phase and multiphase flows, Verlag Sauerlander publishing house, Aarau and Frankfurt am Main, 1971.

The pressure loss PD is determined in the single-phase current by an empty tube by PD = Cw * p / 2 * w2 / D with p [kg / m3] Density of the current medium under operating conditions w [w / s] Speed of the current (volumetric current / cross-sectional area) D [m] Diameter of the tube and Cw [-] Aerodynamic penetration coefficient In the case of a current through a filler body, layers or additional elements, the speed w must be replaced by the effective speed (w /?), as well as the diameter of the tube D by the hydraulic diameter of the dH channel of the filling bodies or elements, for which it becomes valid: PD = Cw * p / 2 * (w /?) 2 * l / dH with dH [m] Hydraulic diameter of the channel? [-] Fraction of the empty tube Cw [-] Coefficient of aerodynamic penetration of the device through which the current passes: with filling.

Often, the specific data concerning the filler bodies dH and? they are part of the technical data for the supply of filler bodies. For a series of llandores bodies, data is indicated in the work "VDI-remeatlas" mentioned above. The fraction of the empty tube can also be determined by means of experiments, for example, by checking the capacity in liters of the reactor before and after filling it with filling bodies. On the other hand, in the case where the hydraulic diameter of the channel is known, it can be calculated from the specific surface F [m2 / m3] of the filler bodies or additional elements (they are generally known or can be determined by the experimental path) according to the simple relationship dH = 4? / F. In general, the aerodynamic penetration coefficient of tubes, elements and filling bodies is described in dependence on the Reynold Re number that provides information on the current state under the selected conditions. In practically all cases the following relationship can be applied for filler bodies, elements, etc .: Cw = K? / Ren + K2 / Rem where n = 1, m = 0 is applied often (main reflection according to S. Ergun, Cehm, Eng.Program 48, (1948), 89), on = 1, m = 0.1 (reflection according to Brauer et al.). Ki, K2 are specific constants referred to filling bodies that are known from the supply data or from the literature (examples are indicated in "VDI-Wármeatlas" and in Brauer et al.). However, they can also be determined experimentally, by operating a tubular reactor with filling bodies with a liquid at different speeds and from the known data and the measured pressure loss Cw in dependence on Re. Finally, the Reynold Re number is defined as Re = w * (p /?) * D for empty tubes or RE = (w /?) * (P /?) * DH for tubes with additional elements or filling bodies,? [Pa * s] refers to the respective viscosity and p [ka / m3] the density of the flowing medium. The pressure loss in biphasic currents (here gaseous-liquid for synthesis gas / catalytic solution) increases in an overproportional manner. Often, the pressure loss of the biphasic current P? G is related according to Lockhart-Martinelli (in Brauer et al.) To the loss of pressure in one of the two phases, for example with the loss of pressure of the liquid phase pure flowing Pi, and with the relation of the pressure loss of the two phases thought as single flowing phases Pi and Pg. For the calculation of pressure losses in biphasic currents, pressures without dimension are frequently applied according to f2 = P? g / P? and X2 = P? / Pg. The other relation f2 = function (X2) was investigated many times. Examples can be found in the following literature: Y. Sato, T. Hirose, F. Takahashi, M. Toda: "Pressure Loss and Liquid Hold Up in Packed Bed Reactor with Cocurrent Gas-Liquid Down Flow"; J. Chem. Eng. Of Japan, Vol. 6 (No. 2), 1973, 147-152; D. Sweeney: "A Correlation for Pressure Drop in Two-Phase Concurrent Flow in Packed Beds"; AIChE-Journal, Vol. 13, 7/1967, 663-669; V. Weekman, J. E. Myers: "Fluid-Flow Characteristics of Concurrent Gas-Liquid Flow in Packed Beds"; AIChE-Journal, Vol. 10 (No. 6), 11/1964, 951-957; R. P. Larkins, R. P. hite, D. Jeffrey: "Two-Phase Concurrent Flow in Packed Beds"; AIChE-Jounal, Vol. 7 (No. 2), 6 &1961, 231-239 or N. Midoux, M. Favier, J.-C. Charpentier: "Flow Pattern, Pressure Loss and Liquid Holdup Data in Gas-Liquid Down-Flow Packed Beds with Foaming and Nonfoaming Liquids "; J. Chem. Eng. Of Japan, Vol. 9 (No. 5), 1976, 350-356.

Often, the ratio proposed by Midoux that has been verified for many gaseous-liquid systems is used for the calculation. In non-foaming systems, for example, f2 is equal to 1 + l / X + 1.14 / X0 54. In many works, this relation named according to Lockart-Martinelli is graphically illustrated. Detailed descriptions are found in many teaching books on procedural technique and publications, including Brauer et al. Therefore, the pressure loss of the biphasic current Pg? results from the pressure loss experimentally determined or estimated, as already explained in the above, of the pure flowing liquid phase Pi, applying: Pgi = f2 * Pi. In the present case of the preparation of aldehydes by the hydroformylation of olefins, the calculation of the pressure loss is still a little more complex. Apart from the synthesis gas phase and a liquid catalytic phase, the presence of a liquid organic phase must also be considered. This problem can be addressed by determining an additional dimensionless pressure f2org = Pg ?? / P? G, so pressure loss is determined as follows: In general, with a reactor length L [m] it is considered that therefore, the loss of pressure of a multiphase current can be calculated by the common methods of the chemical technique of the procedures. The analogue is valid for the dimensionless pressure loss B defined above, that is to say the load factor of the multiphase reactor. The measurement of the load factor without dimension B represents a basic and necessary condition for the method, according to the present invention; B should be greater than or equal to 0.8, preferably greater than or equal to 0.9 or, particularly preferably greater than or equal to 1. In the range of B > o = 0.8, a top-down operated reactor begins to flood. It should be specifically mentioned that in keeping with these conditions, the advantages of the process, according to the present invention, are also achieved if the reactor is operated from the bottom up or in another direction. At any time, the upper linear current densities of the reactor (B> > 1) which are recognized by the rising pressure differential through the reactor are possible, even if the space-time yields justify the energy consumption which increases in the same proportion. Therefore, an upper limit is determined solely on the basis of practical reflections on energy consumption or complications in the separation of the phases after the reaction. It is therefore obvious that, apart from the volumetric currents of the individual phases, that is, the empty tube velocities deduced from these w = V / (pD2 / 4), also the dimensions of the reactor (length L, diameter D) , as well as, in particular, the data of the filling bodies (hydraulic diameter dH, fraction of veicio tube?) have an important meaning. By means of the appropriate selection of the parameters, the procedure can be adapted without major complications to the most different requirements; it is only important that B is complied with > = 0.8, preferably B > = 0.9 and, of particular preference B > = 1. In the case of a slow reaction, for example, the hydraulic diameter of the filled bodies would be selected small or their specific large surfaces, so that the conditions required for B are already met in the case of lower speeds of the current. In this way, sufficient waiting times are obtained for the length of a reactor from a reasonably sized technical point of view. If it is very fast reactions, it is recommended to proceed in inverted form. An additional criterion for carrying out the method, according to the present invention, is the ratio between the mass flow rate of the liquid phase containing the catalyst M1 and the mass flow rate of the phase or the dispersed phases M2. In the present case of hydroformylation, the mass flow rate of the catalytic phase Mi is considerably greater than the mass flow rate of the dispersed phases, ie the organic olefinic phase M2a and the synthesis gas phase M2b. In the process, according to the present invention, the mass ratio Mx / M2 between the continuous phase (Mi) and the dispersed phases (M2) can be greater than 2, preferably M1 / M2 >; 10. Current relationships with M? / M2 > 100 are definitely possible and often even advantageous. Under condition M? / M2 > 2, the catalytic phase is the continuous phase while the dispersed phases are distributed in fine bubbles or fine drops. According to the process according to the present invention, it is possible that at least one educt (olefin) is dispersed by the energy introduced into the tubular reactor by the continuous phase (catalyst). This leads to a distribution of at least one educt in bubbles or fine droplets within the continuous catalytic phase. Also the above can be estimated by common methods in technical engineering. Principles with characteristic numbers without dimensions are appropriate as ds / dH = k * Regi (g? I) m * gi (g? I) n where ds are the diameter of drops or bubbles according to Sauter (in Brauer et al.), DH hydraulic diameter of the filling bodies, Reg? G ??) Reynold number of multiphase current = Wgi (gii) * (p? / ??) * (dH /?), Wg? Gii) Weber number of the multiphase stream Wgi (gii) 2 * (p? / Sgl) * (dH /?) 2, k, m, n empirical constants (known or determined by tests), w empty tube velocities [m / s] = V / (pD2 / 4), V volumetric current under operating conditions [m3 / s], p density under conditions operating [kg / m3],? viscosity under operating conditions [Pa * s] and? interfacial tension under operating conditions [N / m] and indices 1 (liquid phase), g (gas phase), gl (gaseous / liquid biphasic stream) and gil (gaseous / liquid / liquid three-phase stream). It seems that in the case of structured packages such as Sulzer-SMV or narrow tubes as additional elements it is not reasonable that a diameter of bubbles or drops calculated ds is greater than the diameter of the channel. However, the above is not valid for permeable packages and filler bodies such as, for example, metal wire rings or metallic fabrics (so-called dropper packages). In the process, according to the present invention, calculated droplet diameters that are at least equal to or less than the hydraulic diameter of the channel can be used: ds / d < = 1, preferably 0.9. Finally, from the calculated diameter of the droplets, a transformation plane of substances can be calculated according to As = 6fgds [m2 / m3]. For the fraction fg of the dispersed phase (in the case of hydroformylation the synthesis gas and / or the organic phase are dispersed) it can be indicated with the empty tube velocities of the phases fg-W9 / Wgl.

The waiting time t of the phases flowing through the reactor can be calculated approximately according to L *? / W? G. In the method, according to the present invention, the waiting time t, usually, is much less than one hour and may be within the range of minutes or even less. However, with this completely unaccustomed method - high catalyst throughput in the reactor, a small compared part of the reactant in the reaction mass, due to the above a very short waiting time - surprisingly high space-time yields are achieved. . Alternatively, it is possible to work with considerably lower temperatures than usual, obtaining the same space-time yields, since the increase in the speed of the reaction, which for example can lead to a minimization of the sequential reactions and, therefore, Therefore, at an improved selectivity, it allows it economically. The method, according to the present invention, can be adjusted very flexibly to the most different requirements. For special requirements, the following embodiments of the method according to the present invention are offered: If the purpose of use requires a very long mixing zone or if rest areas are required, for example for the extraction of substance streams, an arrangement in the form of a cascade of tubular reactors with additional elements or filling bodies. An arrangement of tubular reactors in the form of a cascade or the alternative arrangement of packed or empty sections of the tube is recommended if a particularly small pressure loss is desired. In addition, it is possible to use a parallel arrangement of tubular reactors or the use of a multi-tubular reactor, where the tubes can accept the function of the additional elements. In addition, the reactors can have a multiple gas supply for the length of the reactor if the gas consumption is so high that inconvenient relations between the gas and liquid phases result in the union of these two phases. The particular conditions of the process, in accordance with the present invention, allow other embodiments of the method. Thus, the necessary and high circuit of the catalytic phase, that is to say, the continuous phase can be further used for the operation of a jet nozzle disposed in front of the tubular reactor itself as a compressor of the liquid jet and of the gas. This can be used to pre-mix the two phases thoroughly, as well as to compress the gaseous phase1, which makes this method possible with high pre-pressures in the reactor. It is also offered in the use of gaseous olefins. Finally, if in the inverted form the suction effect is used instead of the compression of the gases, even a gas circuit conduction is possible under simultaneous pre-mixing of the phases. In this way, the energy introduced by the continuous phase containing the catalyst in the tubular reactor can be used for the dispersion of the phase of the educt, that is, of at least one educt. Also the removal of heat in highly exothermic reactions such as the hydroformylation of olefins is not critical in the process according to the present invention. The high flow rate of the catalytic circuit functions as a thermal carrier, so that even in the case of an adiabatic operation of the reactor, only minimal differences in temperature occur and a thermal distribution in the reactor results without temperature peaks. The heat generated can be removed comfortably by a conventional heat exchanger arranged in the external catalytic circuit or it can be used for power generation. For the purpose of an improved heat removal it may be advantageous if the catalytic circuit is operated at higher levels (ie with a higher B value) than is technically necessary, since a lower temperature gradient can be set by means of the catalytic circuit. by the reactor. The method, according to the present invention, compared to the state of the L-act of the technique offers considerable advantages, these being among others: • With comparably lower temperatures high space-time yields can be obtained. • The formation of secondary products is extremely inferior; Values between 1 and 2% by weight and below are possible. • The catalyst is taken care of, deactivation is very small, continuous extraction is not required. In the present case of the preparation of aldehydes by the hydroformylation of olefins, by applying the process, according to the present invention, other advantages arise: • Based on the higher reaction rate, this process can also be economically used for hydroformylation of higher olefins of more than 6 carbon atoms. • In the case of olefins in the form of a gas, the fraction of educt remaining after a partial reaction can be carried back easily by means of a jet nozzle. In the process according to the present invention, the catalytic phase is the continuous phase; a mass ratio between the catalytic phase and the dispersed phase (s) is advantageous, ie the olefin phase (s) at the reactor inlet within a range of 5000/1 to 4/1, preferably from 2000/1 to 50/1. The mass ratio between the catalytic phase and the hydroformylation agent (usually synthesis gas) is 4/1 to 10000/1, preferably 200/1 to 4000/1. The reactants can be added in preheated form, ie, within the range of the reaction temperature, or cold. Based on the high phase re-phase with the catalytic phase, the previous heating can be carried out by the heat that comes from the process. The process, according to the present invention, for the hydroformylation of olefins is preferably carried out within a temperature range of 20 ° C to 250 ° C, particularly preferably within a range of 90 ° C to 150 ° C. C. Here, the total pressure is between 10 bar and 300 bar, preferably between 20 and 150 bar The current can pass through the tubular reactor in parallel current from top to bottom or in the reverse direction. Due to security reasons, the operation above is preferred. The heat of the reaction can be eliminated by different heat exchangers. Here, the heat exchangers do not necessarily have to be close to the reaction space, but may also be located outside the reactor, as appropriate. The individual heat currents depend on the specific heat of the reaction, as well as on the desired temperatures in the reactor and in the processing devices. The heat of the removed reaction can be used very simply, for example, in the process itself, for the heating of a distillation device or for the generation of steam. In the case of the use of olefins in the form of a gas or of an incomplete reaction in a gas / liquid separator container, the mixture leaving the reactor can be degassed. The gas / liquid separation can be carried out under the same pressure that is present at the reactor outlet. The above is particularly advantageous if at least a part of the expansion gas is returned to the reactor. It can also be expanded under low pressures (down to 1 bar).

The separated gaseous stream can be driven back completely or in part to the reactor. This return conduit can be carried out according to known methods, for example, by means of a jet nozzle or mixer, arranged in the current of the catalytic circuit in front of the reactor, or by means of a compressor of the gas in circuit. From energetic reflections preference is given to a jet nozzle or mixer disposed in the current of the catalytic circuit in front of the reactor. The shelf or optionally complete amount of gas can be introduced in cooled or uncooled form to a waste gas recycling system. If a refrigerator is used, the gas condensate generated in the refrigerator can be conducted through a conduit to the gas / liquid separator container. In a liquid / liquid separator container, the degassed liquid mixture is separated mechanically in the catalytic phase and the product phase. This can be carried out in a settling tank of different types or in centrifuges. Due to economic reasons preference is given to the decanting vessels. Waiting times in the separation device are not primarily critical, however, it is advantageous to keep them short. This has the following advantage: The separating device is small and the inversion correspondingly lower. With short waiting times virtually no side reactions occur in the separating vessel. So that the separation of the phases is carried out quickly, the difference between the densities of the two phases must be sufficiently large and their viscosities low. The three parameters are a function of temperature and can be easily determined by orienting tests. In addition, the density and viscosity of the catalyst solution can be varied by selecting the solvent from the catalyst concentration. As a further option, the density and viscosity of the product phase can be modified by the addition of a solvent. The separation of the phases can be carried out within a wide temperature range. The separating temperature can also be higher than the temperature. of the reaction product at the outlet of the reactor. However, due to energy reasons, it is not recommended to apply a temperature higher than the temperature of the liquid in the gas separator. The solidification point of one of the two phases should be considered as the lowest temperature. However, considering the short separation times, as mentioned in the above, temperatures not too low are selected. The product stream can be divided according to known methods, for example, by distillation. The separated catalyst solution, if necessary after removal of a small partial amount and the corresponding substitution with fresh catalyst solution, is conducted back to the reactor. The following examples are intended to explain the present invention in more detail, without limiting its spectrum of application that can be inferred from the claims: Hydroformylation of propene: Example 1 (Comparative example, batch reaction) In a stirrer autoclave were mixed at a temperature of 120 ° C and 50 bar of synthesis gas 290.3 g of TPPTS ligands (triphenylphosphine trisulfonate) in the form of a sodium salt, 31.8 g of propene and a part of 291 g of a solvent, consisting of 20% by weight of ethylene glycol and 80% by weight of water. The hydroformylation reaction was initiated by the addition of 0.531 g of rhodium acetate dissolved in the rest of the solvent. After the complete reaction, which could be detected by means of the synthesis gas adsorption curve, a liquid test of the reaction mixture was taken.

All continuous hydroformylation assays (also those tests with another product other than propene) were carried out with a test kit shown in figure 1. It was used, as long as the description of the example does not indicate another reactor, a reactor with a length of 3 m and a diameter of 17.3 mm (volume: 705 ml), containing static mixing elements of the company Sulzer with a hydraulic diameter of 2 mm. The aqueous catalyst is pumped by pump 1 in the circuit. Olefin (propene) 3 and synthesis gas 4 are added to the catalytic mixture. The multiphase mezcal 5 obtained in this way is pumped through the tubular reactor 6, passing through the mixing nozzle 11. The tubular reactor 6 has static mixing elements. At this point, the intimate mixing of the phases is of particular importance which, with predetermined mixing elements are a function of the Reynold number. The resulting mixture consisting of the product, the reactant which did not react and the catalyst is degassed in the container 8. Most of the gas 9 consisting of olefin (propene), synthesis gas and enriched inert agents is reintroduced into the reactor 6 by means of a gas reflux conduit 10 and with the aid of a mixing nozzle 11. A small part of the gas stream 9 is expelled outwardly by a conduit 12. By a suitable cooling 13 and the reflux of the Over-critical propene reduces the outgoing substance 14 to enriched inert agents and small amounts of synthesis gas that did not react. By means of this arrangement, the yield of the olefin is practically not limited by the eclusion of the inert agents outwards. The liquid stream 15 that occurs after degassing inside the container is led to a phase separator vessel 16. Here, the catalytic aqueous phase 2 is separated and conducted back to the circuit. The heat of the reaction can be extracted elegantly by an external heat exchanger 17. Examples 2 to 5 (comparative examples, continuous process without solvent mixture) These examples represent comparative tests in the described continuous equipment, in order to emphasize the advantages presented by the present invention in relation to the space-time yield compared with a purely solvent aqueous. For these examples, the conduit 10 for refluxing the gas was closed. Water was used as the solvent for the catalyst. A stream with a catalytic loading of 400 kg / h at a temperature of 120 ° C was passed through the reactor with a volume of 705 ml. The reactor pressure was 50 bar. The rhodium concentration was 800 ppm related to the solvent phase. As Ligand, TSTPP was used in the form of its sodium salt (NaTSTPP); the P / Rh ratio was 60. For example 3 the reaction conditions of example 2 were adjusted with the variation that the reaction temperature was 130 ° C. For test 4 the reaction conditions of example 2 were adjusted with the variation that the reaction pressure was 70 bar. For example 5 the reaction conditions of example 2 were adjusted with the variation that the catalytic loading of the reactor was 300 kg / h. The molar flow rates of the introduced educts, as well as the products are indicated in the table in mol / h.

Examples 6 to 11 (according to the present invention) These examples describe the use, in accordance with the present invention, of solvent mixtures in the continuous equipment described using the water / ethylene glycol example. The measured data emphasize the increased space-time performance. Here, no noteworthy formation of dioxolanes was detected contrasting the agitated systems. In Example 6 the reaction conditions of Example 5 were adjusted with the variation that was used as solvent a mixture of water and ethylene glycol (20% by weight); in Example 7 the concentration of ethylene glycol was increased to 40% by weight. In Example 8 the reaction conditions of Example 7 were adjusted with the variation that the catalytic loading of the reactor was 400 kg / h and the reaction temperature of 130 ° C. In Example 9 the reaction conditions of Example 8 were adjusted with the variation that the reaction temperature was 120 ° C and the reaction pressure of 70 bar. In Example 10, the reaction conditions of Example 9 were adjusted with the variation that the reaction temperature was 90 ° C. This trial documents the possibility of hydroformylation of propene in a multiphase system also at lower temperatures and even with high space-time yields. In Example 11 the reaction conditions of Example 10 were adjusted with the variation that the reaction pressure was 50 bar and the rhodium concentration of 200 ppm relative to the solvent. This essay documents the possibility of carrying out the hydroformylation of propene in a multiphase system also with low concentrations of rhodium and even with high space-time yields. The introduced flows of the educts, as well as the products are indicated in the table in mol / h.

Selective hydroformylation of 1-butene Another application of the process, according to the present invention, is the selective hydroformylation of 1-butene from a refined product mixture I. This mixture consisting of olefins of 4 carbon atoms and paraffins of 4 Carbon atoms contain 1-butene fractions of 26 to 29% by weight. A mixture of water and ethylene glycol (50/50% by weight) was used as catalyst solvent. Example 12 (comparative example, batch process) This example represents a comparative test in an agitator autoclave, in order to emphasize the advantages of the present invention in relation to the quality of the product and in comparison with common agitator reactors. In a stirring autoclave 29.81 g of NaTSTPP, 86.2 g of 1-butene, 13.5 g of isobutane and 67.7 g of a solvent of which 30% by weight is ethylene glycol and 70% by weight is water at a temperature of 105 ° were mixed. C and 30 bar of synthesis gas. The hydroformylation reaction was initiated by the addition of 2.09 g of a rhodium acetate solution containing 3.7% Rh (30% by weight of ethylene glycol and 70% by weight of water). After the complete reaction that could be detected by the synthesis gas absorption curve, a sample of the reaction mixture was taken. For the following relationship, isobutane was calculated as an internal standard.

The continuous hydroformylation of 1-butene in Examples 13 to 26 was performed as the hydroformylation of propene, only that the gas reflux conduit 10 was closed. Example 13: This example describes the use according to the present invention of a solvent mixture of 50% by weight of water and 50% by weight of ethylene glycol in the continuous equipment described in the above. The significantly reduced formation of dioxolanes compared to the agitated systems is documented from the measured data. A catalytic load of 400 kg / hr was passed through the 3 m long reactor at a temperature of 115 ° C. For the above, 600 Nl / h of synthesis gas and 3 kg / h of refined product I were added. The reaction pressure was 50 bar. The following data were determined: Example 14: The reaction conditions of example 13 were adjusted with the variation that the catalytic loading of the reactor was 200 kg / h.

Example 15: The reaction conditions of Example 13 were adjusted with the variation that the catalyst load in the reactor was 100 kg / hr and the reaction temperature 85 ° C.

Example 15: The reaction conditions of Example 13 were adjusted with the variation that the catalyst load in the reactor was 250 kg / h and the reaction temperature 85 ° C.

Example 17: The reaction conditions of Example 13 were adjusted with the variation that the reaction temperature was 90 ° C.

Example 18: The reaction conditions of Example 13 were adjusted with the variation that the reaction temperature was 70 ° C.

Example 19: The reaction conditions of example 13 were adjusted with the variation that the reaction temperature was 80 ° C.

Example 20: The reaction conditions of Example 13 were adjusted with the variation that the reaction temperature was 80 ° C.

Example 21: The reaction conditions of example 13 were adjusted with the variation that the reaction temperature was 95 ° C.

Example 22: The reaction conditions of Example 13 were adjusted with the variation that the reaction temperature was 105 ° C.

Example 23: The reaction conditions of Example 13 were adjusted with the variation that the reaction temperature was 115 ° C.

Example 24: The reaction conditions of Example 13 were adjusted with the variation that the reaction pressure was 33 bar.

Example 25: The reaction conditions of example 13 were adjusted with the variation that the reaction pressure was 43 bar.

Example 26: The reaction conditions of example 13 were adjusted with the variation that the length of the reactor was 1 m.

Example 27: The reaction conditions of example 13 were adjusted with the variation that the length of the reactor was 2 m (with the same filling and the same diameter).

Hydroformylation of 1-decene In Examples 28 and 29, a current of 400 kg / h of catalytic solution was passed through the reactor of the test equipment. The reaction temperature was 125 ° C and the reaction pressure was 70 bar. The rhodium concentration was 800 ppm relative to the catalytic phase. As a ligand, TSTPP was used in the form of its sodium salt. Example 28 (comparative example, continuous process without solvent mixture) Water was used as the solvent for the catalyst. The pH value was 4.5. The P / Rh ratio in the catalyst was 5. The flow rates of the educts, as well as the products, are indicated in the table in mol / h.

Example 29 (according to the present invention) The example describes the use, according to the present invention, of a solvent mixture of 50% by weight of water and 50% by weight of ethylene glycol for the hydroformylation of 1-decene, with In order to document the great advantages of the present invention with respect to the space-time performance compared to the water solvent.

As a solvent for the catalyst, a mixture of water and ethylene glycol (1: 1) was used. The pH value was 7.3. The P / Rh ratio in the catalyst was 60. The introduced flows of the educts and the products are indicated in the table in mol / h.

Claims (14)

1. Claims 1. A process for the hydroformylation of one or more olefins of 2 to 25 carbon atoms, respectively, by multiphase reactions in a tubular reactor, characterized in that a) the continuous phase contains the catalyst, b) the continuous phase contains a mixture solvent, c) the dispersed phase contains at least one olefin and d) the load factor of the tubular reactor is equal to or greater than 0.8.
2. 2. A method according to claim 1, characterized in that the solvent mixture has a dielectric constant of 50 to 78.
3. 3. A process according to claim 1 or 2, characterized in that the solvent mixture consists of water and an organic solvent mixable with water and contains at least two oxygen atoms.
4. 4. A method according to any of claims 1 to 3, characterized in that the catalyst contains a metal of the eighth secondary group of the periodic system of the elements.
5. 5. A process, according to any of claims 1 to 4, characterized in that the catalyst contains rhodium.
6. 6. A process according to any of claims 1 to 5, characterized in that water-soluble rhodium compounds are used as a catalyst.
7. 7. A method according to any of claims 1 to 6, characterized in that the load factor B is equal to or greater than 0.9.
8. 8. A method, according to any of claims 1 to 6, characterized in that the load factor B is equal to or greater than 1.0.
9. A method, according to any of claims 1 to 8, characterized in that the mass ratio between the continuous phase and the phase or dispersed phases is greater than 2.
10. 10. A method, according to any of the claims 1 to 9, characterized in that in front of the reactor the continuous phase drives a jet nozzle.
11. 11. A process according to any of claims 1 to 10, characterized in that at least one educt is dispersed by the energy introduced by the continuous phase into the tubular reactor.
12. 12. The use of the aldehydes prepared according to claims 1 to 11 for the preparation of alcohols.
13. 13. The use of the aldehydes prepared according to claims 1 to 11 in aldol condensations.
14. 14. The use of the aldehydes prepared according to claims 1 to 11 for the preparation of carboxylic acids. ¿; .Ü-J--.