METHOD FOR PROCUREMENT OF ALCOHOLS FROM ALDEHYDES

Description

[0001] The present invention relates to a process for the production of saturated C n - and C 2n - alcohols, wherein the ratio of C n - to C 2n - alcohols is controlled by the diversion of partial streams in individual process steps.

[0002] Alcohols play an important role in industry. They are frequently used as intermediates in the production of lubricating oils, greases, plasticizers, pharmaceuticals, cosmetics, and flavorings. Alcohols are also used directly as solvents, antifreeze, and fuel additives.

[0003] Plasticizers are used in large quantities to modify the thermoplastic properties of a wide variety of industrially important products, such as plastics, but also paints, coatings, sealants, etc. An important class of plasticizers are the ester plasticizers, which include, among others, phthalic acid esters, trimellitic acid esters, and phosphoric acid esters. The alcohols used to produce ester plasticizers are generally referred to as plasticizer alcohols. For the production of ester plasticizers with good application properties, plasticizer alcohols with approximately 5 to 12 carbon atoms are required.

[0004] Due to the ongoing phthalate debate surrounding plasticizers, the demand for new phthalate-free plasticizers is increasing. However, it is crucial that the respective plasticizers meet strict requirements regarding their properties, depending on their intended applications. Examples of such requirements include viscosity and volatility. Controlling the essential properties of plasticizers depends less on the esterification reaction typically used in their production and more on the raw materials employed, particularly the alcohols used. Key factors here include the number of carbon atoms in the alcohols used and their isomer distribution. Alcohols with 4, 5, or 6 carbon atoms are particularly suitable. At the same time, however, it is also necessary to produce the aforementioned C8, C10, and C12 alcohols.

[0005] One of the best-known routes to alcohols is the hydroformylation reaction, in which alkenes are converted to aldehydes, which are then subjected to hydrogenation to obtain the corresponding alcohols (Cₙn alcohols). An exception to this is the hydroformylation of propene and unbranched butenes. Here, the resulting aldehydes usually undergo a further reaction step, aldolization, to obtain long-chain unsaturated aldehydes. These are then also subjected to hydrogenation, and the resulting longer-chain alcohols (C₂n alcohols) are primarily used in the production of phthalate-containing plasticizers.

[0006] For example, US 2014 / 350307 describes the production of alcohols from olefins by hydroformylation, aldol condensation and subsequent hydrogenation of the aldol condensate mixture.

[0007] US 6455743 describes the production of alcohols by hydroformylation of an olefin, aldol condensation of a portion of the aldehyde, and subsequent hydrogenation reaction.

[0008] Prior art processes have the disadvantage that the product compositions, especially of the final alcohols, depend on the starting material streams used, particularly the olefins. Precise control of the composition of the resulting alcohols is therefore not possible, especially with regard to fluctuations in the starting material compositions.

[0009] This presents the challenge of controlling, on the one hand, the isomer distribution of the alcohols to be produced, and on the other hand, in particular, the ratio of C n - to C 2n - alcohols.

[0010] The object of the present invention is to provide a method that meets the above-mentioned requirements and allows the targeted control of the composition of the alcohols obtained.

[0011] The object of the present invention is solved by a process for the production of saturated C n - and C 2n - alcohols, wherein the ratio of C n - to C 2n - alcohols is determined by targeted control of the reactant streams of aldolization and hydrogenation.

[0012] Accordingly, the present invention relates to a process for the production of C n - and C 2n - alcohols with n = 5 comprising the process steps a) Providing a mixture of isomeric Cn-aldehydes with n = 5, wherein the proportion of unbranched aldehydes is at least 40 wt%, based on the Cn-aldehydes with n = 5 provided, and splitting the mixture into a first partial stream and a second partial stream via a split control; b) Carrying out an aldol condensation of the aldehydes contained in the first partial stream to obtain a mixture of Cn- and α,β-unsaturated C2n-aldehydes; c) Mixing the mixture of Cn- and α,β-unsaturated C2n-aldehydes obtained in b) with the second partial stream provided in a); d) Hydrogenating the mixture of Cn- and α,β-unsaturated C2n-aldehydes and isomeric Cn-aldehydes obtained in c) with n = 5 with hydrogen to obtain a mixture of saturated C n and C 2n alcohols and e) separation of the mixture of C n and C 2n alcohols by at least one two-column system or by at least one dividing wall column, characterized in that the isomer ratio of the obtained C n -Alcohols are adjusted by the ratio of the first and second partial streams generated in step a) and by the aldol condensation carried out in step b).

[0013] The process chain according to the invention surprisingly allows the production of Cn and C2n alcohols that can be used directly for the production of plasticizers. In particular, it has been found that the isomer distribution, for example of the Cn alcohols, is such that after esterification they lead to plasticizers exhibiting an advantageous range of properties, and the water content of the Cn alcohols is so low that they can be used without further pretreatment in anhydride esterification reactions or transesterifications.

[0014] A key feature of the present invention is that the isomer ratio of the resulting Cn alcohols can be adjusted. According to the invention, this is achieved, firstly, by adjusting the ratio of the first to the second partial stream during the splitting process in step a), and secondly, by the aldol condensation carried out in step b). During the splitting into a first and a second partial stream in step a), the amount of Cn aldehyde bypassing the aldol condenser can be actively controlled. This provides a significant control parameter for the isomer distribution of the resulting Cn alcohols. The parameters of the aldol condensation are another important control parameter, as they allow for the control of the yield of C2n alcohols and thus automatically the proportion of unreacted Cn aldehydes (and therefore Cn alcohols).Furthermore, the isomer ratios can also be controlled by selectivity effects during aldol condensation. Thus, the aldehydes used generally exhibit different reaction rates in the aldolization. For example, n-pentanal reacts significantly faster than 2-methylbutanal.

[0015] The basic sequence of the method according to the invention is in Fig. 1The process is illustrated. In the first step, a stream of isomeric Cn-aldehydes with n = 5 is provided (1) and divided into a first partial stream (1a) and a second partial stream (1b). The first partial stream (1a) containing isomeric Cn-aldehydes with n = 5 is introduced into a reactor to carry out an aldol condensation, yielding a mixture of Cn- and α,β-unsaturated C2n-aldehydes. The product stream obtained from the aldol condensation, containing a mixture of C n and α,β-unsaturated C 2n aldehydes, is mixed with the partial stream (1b) obtained in the first step, containing isomeric C n aldehydes with n = 5, and transferred as reactant stream (2) to a further reactor for the purpose of hydrogenating the mixture of C n and α,β-unsaturated C 2n aldehydes and isomeric C n aldehydes with n = 5 u (reactant stream (2)) with hydrogen to obtain a mixture of saturated C n and C 2n alcohols.

[0016] Thermal separation of this mixture yields the respective fractions with C n and C 2n alcohols as the main components.

[0017] In principle, all Cn-aldehydes known to those skilled in the art with n = 5, either in pure form or in the form of mixtures, are suitable for use in processes according to the present invention, wherein a basic requirement is that the proportion of unbranched aldehydes is at least 40 wt.%, based on the Cn-aldehydes used with n = 4, 5 and 6.

[0018] In the process according to the invention, aldehyde streams of the same chain length are used in particular. The proportion of unbranched aldehydes is at least 40 wt.%, based on the Cn aldehydes used with n = 5. In particular, the proportion of unbranched aldehydes is 40 to 99.5 wt.%, and most preferably the proportion of unbranched aldehydes is 95 to 99.5 wt.%, based on the Cn aldehydes used with n = 5.

[0019] In many large-scale industrial processes, such aldehydes or mixtures thereof are obtained from the corresponding olefins by hydroformylation.

[0020] Accordingly, in a preferred embodiment of the present invention, the C n aldehydes with n = 5 provided in step a) are produced by hydroformylation of isomeric olefins with 3 to 5 carbon atoms with synthesis gas in the presence of a hydroformylation catalyst to form the aldehydes mentioned.

[0021] Corresponding hydroformulation processes are known to those skilled in the art and are described, for example, in Hydroformylation Fundamentals, Processes and Applications in Organic Synthesis Volume 1 & 2 Edition 1, Franke, Börner, Willey VCH Verlag GmbH & Co, Weinheim.

[0022] Typically, rhodium or cobalt catalysts are used, with or without complex-stabilizing additives such as organic phosphines or phosphites. Temperatures and pressures can vary widely depending on the catalyst or olefin. A description of the hydroformylation of olefins can be found, for example, in J. Falbe, *New Syntheses with Carbon Monoxide*, Springer-Verlag, Heidelberg-New York, 1980, page 99 ff., and in Kirk-Othmer, *Encyclopedia of Chemical Technology*, Volume 17, 4th edition, John Wiley & Sons, pages 902 to 919 (1996).

[0023] The hydroformylation reaction mixtures are advantageously first freed from the catalyst before use in the process according to the invention. If a cobalt catalyst has been used, this can be done by depressurization, oxidation of the cobalt carbonyl compounds remaining in the hydroformylation mixture in the presence of water or aqueous acid, and separation of the aqueous phase. Cobalt removal processes are well known; see, e.g., B.J. Falbe, op. cit., Kirk-Othmer, op. cit., 164, 175, BASF process.

[0024] When a rhodium compound is used as a hydroformylation catalyst, it can be separated as a distillation residue, for example, by thin-film evaporation.

[0025] In the preferred embodiment of the present invention, the hydroformylation is carried out according to the method described in WO 2017 / 080690.

[0026] In hydroformylation, a catalyst system is used that contains rhodium as the central atom and is complexed with the ligand (1):

[0027] The IUPAC name of ligand (1) is 3,3'-di-tert-butyl-5,5'-dimethoxy-[1,1'-biphenyl]-2,2'-diyltetrakis(2,4-dimethylphenyl)bis(phosphite).

[0028] Hydroformylation is carried out particularly at a temperature between 120 °C and 140 °C. The pressure is preferably between 15 x 10⁵ Pa and 25 x 10⁵ Pa.

[0029] To extend the service life, the hydroformylation is carried out in the presence of an organic amine of formula (2), wherein Ra, Rb, Rc, Rd, Re and Rf represent identical or different hydrocarbon residues, which may also be linked together. The organic amine preferably comprises at least one 2,2,6,6-tetramethylpiperidine unit. Specifically, the organic amine may be a sebacic acid di-4-(2,2,6,6-tetramethylpiperidinyl) ester.

[0030] It is recommended to adjust the rhodium concentration in the first reaction mixture to between 1 wt. ppm and 1000 wt. ppm. The ligand / rhodium ratio should be between 1:1 and 100:1, with no other ligand being required as part of the homogeneous catalyst system besides the organophosphorus compound according to formula (1). In industrial operation, it cannot be ruled out that, due to impurities, organophosphorus compounds other than 3,3'-di-tert-butyl-5,5'-dimethoxy-[1,1'-biphenyl]-2,2'-diyltetrakis(2,4-dimethylphenyl)bis(phosphite) complex with the rhodium as part of the catalyst system. However, such impurities are not to be taken into account with the specified ligand / rhodium ratio. This specification refers solely to ligand (1), and no other ligand is intended to be included.

[0031] The isomeric C n aldehydes with n = 5 used in the process according to the invention are divided into a first partial stream (1a) and a second partial stream (1b) according to step a).

[0032] The flow is divided via a split control system. This is achieved through corresponding flow control of the two individual flows, 1a and 1b. This control system is implemented according to the state of the art, for example, through the use of control valves and a suitable process control system.

[0033] As already stated, the control of the ratio of the first to the second partial stream is an essential control element for adjusting the isomer ratios of the resulting C n alcohols.

[0034] The ratios between 1a and 1b are adjusted so that, according to the following approach, the desired amount of C 2n alcohol and C n alcohol of the desired isomer distribution is produced: Massenstrom C 2 n Alkohol 5 = Feed * Split * Xn_Aldol * Feed_w_n + Xi_Aldol * Feed_w_i * 2 * M _ Cn _ al − 18 / 2 * M _ Cn _ al + 2 * X _ Hydr * X _ Dest Massenstrom Cn Alkohol 4 = Feed * Split * 1 − Xn _ Aldol * Feed _ w _ n + Feed * 1 − Split * Feed_w_n / M_Cn_al * M_Cn_ol * X_n_Hydr * X_n_Dest + Feed * Split * 1 − Xi_Aldol * Feed _ w _ i + Feed * 1 − Split * Feed _ w _ i / M _ Cn _ al * M _ Cn _ ol * X _ i _ Hydr * X _ i _ Dest Massenstrom Umgang 1 b = Feed * 1 − Split Anteil lineare Alkohole 4 = Feed * Split * 1 − Xn_Aldol * Feed_w_n + Feed * 1 − Split * Feed _ w _ n / M _ Cn _ al * M _ Cn _ ol * X_n_Hydr * X_n_Dest / Feed * Split * 1 − Xn_Aldol * Feed_w_n + Feed * 1 − Split * Feed _ w _ n / M _ Cn _ al * M _ Cn _ ol * X _ n _ Hydr * X _ n _ Dest + Feed * Split * 1 − Xi_Aldol * Feed_w_i + Feed * 1 − Split * Feed_w_i / M_Cn_al * M_Cn_ol * X_i_Hydr * X_i_Dest Feed = Feed (1) Split = Percentage (1a) based on the feed (1) Feed_w_n = Percentage concentration of linear aldehydes Cn in the feed (1) Feed_w_i = Percentage concentration of branched aldehydes Cn in the feed (1) Xn_Aldol = Yield of C2n aldolization product based on the linear aldehydes Cn in the aldolization (B) Xi_Aldol = Yield of C2n aldolization product based on the branched aldehydes Cn of the aldolization (B) M_Cn_al = Molar mass of the aldehydes Cn X_Hydr = Yield of the hydrogenated C2n aldolization products to the corresponding alcohols X_Dest = Yield of the C2n alcohols in the distillation X_n_Hydr = Yield of the hydrogenated linear Cn aldehydes to the corresponding alcohols X_i_Hydr = Yield of the hydrogenated branched Cn aldehydes to the corresponding alcohols X_n_Dest = yield of the linear Cn alcohols in the distillation X_i_Dest = yield of the branched Cn alcohols in the distillation

[0035] In a further embodiment, the isomer distribution of the Cn alcohols can be shifted within limits by using thermal separation (separation according to step e)). The following applies to controlling the quantity, referring only to the purified mass flows of Cn alcohol: M_Dest = − 1 * M_Zulauf * Prod_w_n − Zulauf_w_n / Dest_w_n − Prod_w_n

[0036] The following limits apply: Dest_w_n < Zulauf_w_n < Prod _ w _ n M_Zulauf = mass flow rate of the Cn alcohols Prod_w_n = percentage concentration of the linear aldehydes Cn product ((4) in Figure (2) &3)) Inlet_w_n = percentage concentration of linear aldehydes Cn inlet ((8) in Figure (2) &3)) Dest_w_n = percent concentration of linear aldehydes Cn distillate ((10) in Figure (2) &3)) M_Dest = mass flow rate of Cn alcohols to be discharged ((10) in Figure (2) &3))

[0037] All positions refer only to the alcohol components; minor components are not considered and are excluded from the mass flows.

[0038] To monitor the isomer ratios of the Cn alcohols, the use of appropriate analytical techniques is advantageous. Within the scope of the present invention, the use of gas chromatographic methods is particularly preferred. For this purpose, a gas chromatograph of type 7890A from Agilent with a flame ionization detector (FID) is especially suitable. HP-5 columns (5% phenyl methyl siloxane) with nitrogen as the carrier gas are preferably used. However, suitable spectroscopic inline analytical methods are also conceivable.

[0039] According to the invention, the isomeric Cn-aldehydes with n = 5 are converted in partial stream (1a) by aldol condensation, particularly in the presence of sodium hydroxide, into a mixture of Cn- and α,β-unsaturated C2n-aldehydes. Corresponding processes are known, for example, from DE19957522A1, DE102009045139A1, and DE102009001594A1.

[0040] In particular, the aldol condensation preferably takes place in a tubular reactor which contains at least one mixing module that disperses the reactant aldehyde in droplets with an average diameter (Sauter diameter) of 0.2 mm to 2 mm in the continuous catalyst phase (process lye) which consists of sodium hydroxide and sodium salts of carboxylic acids and has a sodium content of 0.6 to 1.75 wt% and a pH value in the range of 12.5 to 13.5.

[0041] In the process according to the invention, sodium hydroxide is used to form the process lye. The sodium hydroxide, together with the return lye, forms the process lye. Besides sodium hydroxide, the return lye contains sodium salts of carboxylic acids, mainly pentanoic acids. The carboxylic acid salts are essentially formed by a Cannizzaro reaction.

[0042] In the process according to the invention, the sodium content of the process liquor at the reactor inlet is between 0.60 and 1.75 wt%, particularly between 1.1 and 1.20 wt%. To adjust the sodium concentration of the process liquor, sodium hydroxide solution with a concentration greater than 2.5 wt% is fed into the return liquor. To minimize the amount of water introduced into the reaction system, sodium hydroxide solution with a higher concentration is preferably used. In the process according to the invention, sodium hydroxide solution in the concentration range of 5 to 30 wt%, for example 10 wt%, is preferably used.

[0043] The process according to the invention is carried out in a tubular reactor comprising at least one mixing module, preferably several mixing modules. In particular, the number of mixing modules is 1 to 30, most particularly 10 to 20.

[0044] A mixing module is understood as a static mixer, i.e., a passive component that has no direct energy requirement of its own.

[0045] The tubular reactor consists of a tube, preferably oriented vertically. Flow can occur from bottom to top or vice versa. A technical reactor can also consist of several parallel tubes connected by U-tubes.

[0046] Preferably, a mixing module is located at the reactor inlet. Empty spaces are located between the mixing modules. The volume fraction outside the mixing module(s) of the total reactor volume is 20 to 80%, particularly 30 to 60%. The mixing modules can be spaced equally or at different distances from each other. Preferably, the distance between the mixing modules decreases in the flow direction. Depending on the intended flow velocity, the phase ratio between the reactant and catalyst phases, the reaction progress, and the mixer type, the distances between the mixing modules are 0.2 to five times the mixing module length, particularly 0.5 to twice the mixing module length.

[0047] The mixing module consists of a static mixer or an arrangement of several, preferably two, static mixers.

[0048] If the mixer module consists of two identical static mixers, these are preferably arranged rotated about the longitudinal axis of the reactor, in particular rotated by an angle of 45° to 90°. Preferably, mixing elements in the mixer module are arranged with a distance of two pipe diameters.

[0049] A mixing module can also consist of static mixers of different designs. It can be advantageous for a mixing module consisting of two static mixers to have the first one with a smaller hydraulic diameter than the second. This ensures that the first static mixer produces the smallest possible droplets, while the second static mixer, with its larger hydraulic diameter, prevents the coalescence of the droplet swarm.

[0050] The hydraulic diameter of the mixing elements of the mixer modules preferably decreases with the flow direction.

[0051] The mixer modules in the reactor can be the same or different, i.e., they can be of the same or different design.

[0052] Any static mixer capable of dispersing the organic phase in the catalyst phase into droplets with an average Sauter diameter in the range of 0.2 to 2.0 mm under the intended reaction conditions can be used as mixing elements.

[0053] In the inventive method, mixing elements suitable for the dispersion of two immiscible, low-viscosity liquids, such as those commercially available, can be used as static mixers.

[0054] According to the invention, the aldol condensation of the C n -aldehydes is carried out in the temperature range of 100 to 150 °C, in particular in the range of 110 to 140 °C, and especially in the range of 120 to 140 °C.

[0055] The conversion can be carried out isothermally, adiabatically, or polytropically within the specified temperature ranges. For example, the reactor inlet temperature can be 120 °C and the reactor outlet temperature 140 °C.

[0056] The reaction pressure in the reactor is high enough to ensure that both the process solution and the organic substances (reactants and products) are present in liquid phase. The pressure is in the range of 0.2 to 1.0 MPa, preferably in the range of 0.3 to 0.5 MPa.

[0057] In the process according to the invention, the quantity ratio [kg / kg] of process liquor to reactant at the reactor inlet is in the range of 5 to 40, in particular in the range of 10 to 15.

[0058] The average empty tube velocity of the mixture of reactant and process liquor (assuming the same flow velocity of both phases) in the technical reactor is in the range of 0.5 to 4 m / s, particularly in the range of 1 to 2.5 m / s.

[0059] The average residence time of the reaction mixture in the reactor is 40 to 360 s, in particular 60 to 180 s.

[0060] In the process according to the invention, the droplets of the organic phase dispersed in the process solution have an average Sauter diameter of 0.2 to 2 mm, in particular of 0.6 to 1.3 mm, after leaving a mixer module.

[0061] The stress factor is in the range of 0.2 to 0.8.

[0062] The obtained aldehydes can optionally be processed before further use in the process according to the invention.

[0063] One method is to cool the reaction discharge and separate the organic phase from the alkaline phase. Phase separation preferably takes place in the temperature range of 60 to 130 °C, particularly in the range of 70 to 120 °C, and especially in the range of 90 to 110 °C. Separation times range from 3 to 10 minutes, depending on the chosen temperature. At temperatures above 90 °C, the separation time is less than 8 minutes. The separation time is defined as the time after which the organic product phase is clear and free of traces of heterogeneous water.

[0064] To separate the heavy, aqueous phase from the light, organic phase, separators can be used that enable phase separation solely through the use of gravity. These so-called gravity separators can also be designed with internals to promote coalescence and increase separation efficiency. The use of internals accelerates the coalescence and sedimentation process. Examples of coalescence aids include plates, packings, fabric packings, or fiber bed separators. Gravity separators can be designed as horizontal or vertical tanks.

[0065] As an alternative to gravity separators, separators based on the principle of centrifuges can also be used for liquid-liquid separation. Centrifugal forces in a rotating drum separate the heavier phase.

[0066] To separate the heavy, aqueous phase, gravity separators are preferably used in the inventive method, preferably gravity separators designed as horizontal containers with internals.

[0067] Part of the separated alkaline phase is removed to separate the reaction water, while the other part is returned to the reactor. The removed stream also separates some of the carboxylic acids (as sodium salts) and sodium hydroxide formed as byproducts. This stream can be fed into a wastewater treatment plant. However, it is also possible to process this stream and partially reintegrate it into the process, as described, for example, in DE 198 49 922 and DE 198 49 924.

[0068] If the organic phase contains, in addition to the Cn- and α,β-unsaturated C2n-aldehydes, other byproducts such as carboxylic acid salts, sodium hydroxide, and dissolved water, traces of bases and some of the carboxylic acid salts can be removed by washing with water. The resulting water extract can be used to prepare the fresh lye solution.

[0069] As previously explained, the isomer ratios of the alcohols and other substances produced according to the invention are controlled by aldol condensation. In particular, according to the invention, the degree of conversion of the aldehydes used, and thus the isomeric composition of the Cn aldehydes in the initial stream of the aldolization, can be controlled by adjusting the reaction temperature and the alkali concentration in the aldolization reaction. Controlling the residence time of the aldehydes used is also another effective means of controlling their degree of conversion. This can be achieved, for example, by changing the reaction volume and / or by controlling the feed rate. Alternatively, or in combination with the aforementioned parameters, the residence time and the aldehyde concentration can also be influenced by adding suitable solvents to the feed stream, thereby also controlling the degree of conversion.

[0070] The C n - and α,β-unsaturated C 2n -aldehydes thus obtained are mixed with the second partial stream (1b) provided in a) according to step c) to obtain a reactant stream (2).

[0071] The mixing of the two partial flows is ensured by suitable static or dynamic mixers, as is known to those skilled in the art. Here, the two flows are preferably combined using a static mixer, such as the CompaX™ module from Sulzer. A suitable pump can also serve as a dynamic mixer at this point.

[0072] The reactant stream (2) obtained by mixing contains the Cn and α,β-unsaturated C2n aldehydes obtained in the aldolysis and the originally used isomeric Cn aldehydes with n = 5 in the form of a mixture. This mixture is hydrogenated with hydrogen according to step d) of the process according to the invention to obtain a mixture of saturated Cn and C2n alcohols. The hydrogenation is also carried out according to processes known per se, for example in the temperature range of 170°C to 200°C at a pressure of 15 × 10⁵ Pa to 30 × 10⁵ Pa on a supported catalyst which contains at least nickel and copper as active components, as known, for example, from EP3037400.

[0073] Preferably, the hydrogenation catalyst consists of a support material based on titanium dioxide, zirconium dioxide, aluminum oxide, silicon oxide or their mixed oxides, wherein hydrogenation-active metals, in particular at least one element of the group copper, cobalt, nickel, chromium, are applied to this support material.

[0074] Aluminum oxide, aluminosilicate, silicon dioxide, titanium dioxide, and zirconium dioxide can be used as support precursors. Aluminum oxide, particularly γ-aluminum oxide, is a preferred support precursor. The catalyst can contain one or more of the hydrogenation-active metals. Preferably, the catalyst used according to the invention contains the metals copper, chromium, and nickel. Particularly preferably, the catalyst contains a combination of the three metals copper, chromium, and nickel as the hydrogenation-active metal.

[0075] The total content of hydrogenating metals, based on the reduced catalyst, is in the range of 1 to 40 wt%, in particular in the range of 5 to 25 wt%, calculated as metal.

[0076] The catalysts used according to the invention are advantageously produced in a form that offers low flow resistance during hydrogenation, such as tablets, cylinders, extruded strands, or rings. During catalyst production, the support material is typically formed into a suitable shape. Pre-shaped support material is also commercially available.

[0077] In the process according to the invention, the hydrogenation can be carried out continuously or batchwise on suspended, finely divided or shaped catalysts arranged in a fixed bed. Continuous hydrogenation on a catalyst arranged in a fixed bed, in which the product / reactant phase is mainly in the liquid state under reaction conditions, is preferred.

[0078] If hydrogenation is carried out continuously on a catalyst arranged in a fixed bed, it is advantageous to convert the catalyst into its active form before hydrogenation. This can be achieved by reducing the catalyst with hydrogen-containing gases according to a temperature program. The reduction can optionally be carried out in the presence of a liquid phase passed over the catalyst, as described, for example, in DE 199 33 348.

[0079] The hydrogenation process according to the invention is carried out in the trickle phase or, preferably, in the liquid phase in three-phase reactors in cocurrent flow, wherein the hydrogen is finely dispersed in the liquid reactant / product stream in a manner known per se. In the interest of uniform liquid distribution, improved heat removal, and high space-time yield, the reactors are preferably operated with high liquid loadings of 15 to 120, in particular 25 to 80 m³ per m² cross-section of the empty reactor per hour. If a reactor is operated isothermally and in straight-through mode, the specific catalyst loading (SCR) can assume values ​​between 0.1 and 10 h⁻¹.

[0080] The process according to the invention is carried out with hydrogen at a pressure range of 5 to 100 bar, preferably between 5 and 40 bar, and particularly preferably in the range of 10 to 25 bar. The hydrogenation temperatures are between 120 and 220 °C, particularly between 140 and 190 °C.

[0081] The hydrogen used for hydrogenation can contain inert gases such as methane or nitrogen. Preferably, hydrogen with a purity greater than 98%, and particularly greater than 99%, is used.

[0082] Different process variants can be selected for the process according to the invention. It can be carried out adiabatically or practically isothermally, i.e., with a temperature rise of less than 10 °C, in one or more stages. In the latter case, all reactors, preferably tubular reactors, can be operated adiabatically or practically isothermally, or one or more can be operated adiabatically and the others practically isothermally. Furthermore, it is possible to hydrogenate the carbonyl compounds or mixtures of carbonyl compounds in the presence of water in a straight-through process or with product recycling.

[0083] In addition to the α,β-unsaturated C 2n aldehydes and the C n starting aldehydes not converted in the aldolization, the hydrogenation also converts the C n starting aldehydes contained in the second partial stream (1b) to the corresponding alcohols.

[0084] According to the invention, the hydrogenation is followed by the separation of the mixture of saturated C n and C 2n alcohols in step e), wherein the separation of the C n alcohols and the C 2n alcohols can be carried out via various combinations of classical distillation columns or dividing wall columns or a combination of both column types by at least one two-column system or by at least one dividing wall column.

[0085] In the present invention, a two-column system or a partition column is used to obtain the C n - and C 2n - alcohols as valuable products in both cases.

[0086] The different procedural variants are explained in more detail below.

[0087] In one embodiment of the present invention, at least one two-column system is used. In the first column, alkanes formed during the process (in the case of butenes, mainly nonane) are drawn off overhead together with the C₅n alcohol (for example, 2-methylbutanol) and water. Phase separation takes place there, and the organic phase is returned to the column as reflux. This aceotropic drying process maintains a water content in the bottoms of less than 1000 ppm. In the second column, the C₅n alcohol is drawn off overhead and thus separated from the high-boiling components. By appropriately adjusting the column operation, the isomeric composition of the C₅ alcohols can be further controlled.

[0088] Fig. 2 shows the detailed process of separating the mixture of C n and C 2n alcohols using a two-column system, which is divided into the following sub-steps: I. Distillative separation (G) of the crude product stream from hydrogenation (3) from step d) into a low-boiling stream (6) and a high-boiling stream (15). The low-boiling stream (6) contains as its main product the alcohols of the starting aldehydes, other substances or azeotropes that have a lower boiling point than the alcohols formed from the α,β-unsaturated C₂n aldehydes, and heterogeneous water. The high-boiling stream (15) contains as its main product the alcohols formed from the α,β-unsaturated C₂n aldehydes and high-boiling substances or azeotropes that have a higher boiling point than the C₂n alcohols. II. Separation of the heterogeneous water (7) from the distillate fraction (6) via separators (F). III. Distillative separation (H) of the organic phase (8) from F into a low-boiling stream (9) and a high-boiling stream (12). The low-boiling stream (9) contains substances orAzeotropes that have a lower boiling point than the highest-boiling Cn alcohol and may contain heterogeneous water (11) after condensation. The high-boiling stream mainly contains the Cn alcohols and substances or azeotropes that have a higher boiling point than the highest-boiling Cn alcohol. IV. Separation of the heterogeneous water (11) by phase separation (E). The organic phase is returned to the column (H) as reflux (19). A portion of the resulting organic phase (10) is removed from the system. V. Distillative separation (I) of the high-boiling stream (12) into a low-boiling stream (4) and a high-boiling stream (14). The low-boiling stream (4) mainly contains the Cn alcohol with the desired isomer distribution. The high-boiling stream (14) mainly contains substances or azeotropes that have a higher boiling point than the highest-boiling Cn alcohol. VI.Distillative separation of the high-boiling stream (15) into a low-boiling stream (5) containing the C₂n alcohols and a high-boiling stream (18). The low-boiling stream (5) mainly contains the C₂n alcohol. The high-boiling stream (18) mainly contains substances or azeotropes that have a higher boiling point than the highest-boiling C₂n alcohol.

[0089] The temperature profiles used in the columns are adapted depending on the respective composition of the product stream from the hydrogenation (step b) of the process according to the invention).

[0090] For example, the product stream (3) may have the following composition: 2-Propylheptenal (mass %) 0,9 n-Pentanol (mass %) 2,8 2-Methylbutanol (mass %) 2,8 2-Propylheptanol (mass %) 87,5 2-Propyl-4-Methylhexanol (mass%) 4,2 H₂O 1,7

[0091] With this composition, operation is preferably carried out at a top pressure of 0.2 bar (abs) and a pressure drop of 0.1 bar in column (G), maintaining a bottom temperature of 176 °C and a top temperature of 65 °C, or equivalent values. In column (H), operation is preferably carried out at a top pressure of 1.05 bar (abs) and a pressure drop of 0.03 bar, with a bottom temperature of 137 °C and a top temperature of 94 °C, or equivalent values. In column (I), operation is generally carried out at a top pressure of 1.08 bar (abs) and a pressure drop of 0.03 bar, with a bottom temperature of 178 °C and a top temperature of 134 °C, or equivalent values.

[0092] The column (J) is operated at a pressure of 0.15 bar(abs) and a pressure differential of 0.006 bar, a sump temperature of 175 °C and a top temperature of 158 °C, or corresponding equivalents.

[0093] The temperature profiles mentioned are to be considered exemplary. Further embodiments, particularly in connection with varying feed compositions, are included within the scope of the present invention.

[0094] In another, equally preferred embodiment of the present invention, at least one dividing wall column is used. In a dividing wall column, the separation operations described for the two-column system can be combined in a single column. The alkane / 2-methylbutanol / water mixture, for example, is also obtained overhead, subjected to phase separation analogously, and the organic phase is returned to the column as reflux. The Cn alcohol, meeting specifications, is drawn off via a side draw-off in the area of ​​the dividing wall. The isomer ratio of the Cn alcohol in the side draw-off can be controlled within certain limits by the distillate withdrawal of the organic phase and the proportion of alcohol in the column bottoms. Fig. 3 The process is shown when using a partition column (K) for the separation of the low-boiling stream (6).

[0095] Here, the in Fig. 2The illustrated columns H and I are combined into a partition wall column (K) (see Fig. 3 ). In the dividing wall column (K), at a top pressure of 1.05 bar(abs) and a pressure drop of 0.06 bar, a bottom temperature of 180 °C, a top temperature of 118 °C and a product temperature of 134 °C, or corresponding equivalents, must be maintained.

[0096] Fig. 4 Figure 1 shows another preferred embodiment in which the separation of the crude product stream from hydrogenation d) is also achieved via a dividing wall column. In this case, the C2n alcohol is drawn off via a side outlet in the area of ​​the dividing wall.

[0097] The process according to the invention is advantageously suitable for the production of C n - and C 2n - alcohols from aldehydes or, in the case of upstream hydroformylation, from olefins.

[0098] In the isomeric Cn-aldehydes used, n = 5. As previously explained, such aldehydes can be obtained by hydroformylation of corresponding isomeric olefins. These isomeric olefins are butenes. Using the process according to the invention, these are reacted to form pentanol mixtures and decanol mixtures.

[0099] With regard to plasticizer syntheses, the isomer distributions of the pentanol mixtures are of particular interest. By way of example and as a preferred embodiment, the resulting pentanol mixture preferably contains less than 60 mol% n-pentanol. The minimum n-pentanol content in the mixture of isomeric pentanols is preferably at least 2 mol%, more preferably at least 10 mol%, more preferably more than 20 mol%, and more preferably more than 22.5 mol% or even more than 25 mol%, and further preferably more than 27.5 mol%, 30 mol%, or even more than 35 mol%. In addition to the linear n-pentyl groups, the particularly preferred pentanols contain branched pentyl groups. A branched pentyl group is preferably a methylbutyl group.Accordingly, a pentanol mixture is preferred in which the branched pentyl groups consist of at least 50 mol%, preferably at least 60 mol%, more preferably at least 70 mol%, more preferably at least 80 mol%, most preferably at least 90 mol% and particularly at least 95 mol% methylbutyl groups.

[0100] It is advantageous if the branched isomeric pentyl groups have a large proportion of 2-methylbutyl groups. In a preferred embodiment, therefore, at least 50 mol%, preferably at least 60 mol%, more preferably at least 70 mol%, further preferably at least 80 mol%, particularly preferably at least 90 mol%, and especially at least 95 mol% of the branched isomeric pentyl groups are 2-methylbutyl groups. The preferred pentanol mixtures preferably contain 20 to 95 mol%, more preferably 30 to 85 mol%, and especially 40 to 75 mol% of 2-methylbutyl groups, based on all pentyl groups contained.

[0101] In a particularly preferred embodiment, the pentanol mixture consists of at least 75 mol%, more preferably at least 90 mol% and in particular at least 95 mol% of pentanols which contain - preferably exclusively - 2 methylbutyl and / or linear pentyl groups, wherein the molar ratio of 2 methylbutyl groups to linear pentyl groups within this pentanol mixture is preferably in the range of 99:1 to 40:60, in particular in the range of 70:30 to 40:60.

[0102] The adjustment of the desired product properties, in particular the aforementioned isomer distributions, is achieved in particular by the parameters described in the respective process steps a), b) and e), especially by the sodium hydroxide concentration used in the aldolization and the temperatures mentioned there (see step b)) as well as by the parameters for carrying out the separation of the mixture of C n and C 2n alcohols by the two-column or dividing wall column systems according to the invention.

[0103] Even without further explanation, it is assumed that a person skilled in the art can use the above description to the fullest extent. The preferred embodiments and examples are therefore to be understood merely as descriptive disclosures, and in no way as limiting disclosures.

[0104] The present invention is explained in more detail below with reference to examples. Alternative embodiments of the present invention are available in an analogous manner. Examples:

[0105] The following basic assumptions were used for carrying out the examples and determining the respective parameters, based on the system C n / C 2n with n = 5.

[0106] The conversion rate Xi1 and Xi2 were calculated according to the following calculation rule: Xi 1 = Masse % nPentanal Edukt Eingang − Masse % nPentanal Produkt Ausgang Masse % nPentanal Edukt Eingang Xi 2 = Masse % 2 Methylbutanal Edukt Eingang − Masse % 2 Methylbutanal Produkt Ausgang Masse % 2 Methylbutanl Edukt Eingang

[0107] The calculation of the n-pentanal fraction based on the C5 aldehydes in the product (Sn1) was carried out according to the following calculation rule: S n 1 = Masse % nPentanal Edukt Ausgang Masse % nPentanal Produkt Ausgang + Masse % 2 Methylbutanal Produkt Ausgang

[0108] The calculation of the 2-propylheptenal fraction based on the unsaturated C10 aldehydes in the product (Sn2) was carried out according to the following calculation method: Sn 2 = Masse % 2 Propylheptenal Masse % 2 Propylheptenal + Masse % 4 Propyl 4 Methylhexenal

[0109] The material compositions in the examples were determined by gas chromatography. An Agilent 7890A gas chromatograph with a flame ionization detector (FID) was used for this purpose. An HP-5 column (5% phenyl methyl siloxane) at 325 °C, 30 m x 320 µm x 0.25 µm, was used.

[0110] The following temperature program was used for the analysis: 2 min at 50 °C, then up to 150 °C at 10 °C / min, followed by up to 250 °C at 40 °C / min with a 10 min holding period. Injection volume 0.5 µl. Injector temperature 250 °C. Carrier gas nitrogen at a flow rate of 105 ml / min and a 50:1 split. Example 1: Aldolization of n-pentanal and 2-methylbutanal including condensation to α,β-unsaturated C10 aldehydes in a flow tube

[0111] The α,β-unsaturated C10 aldehydes were produced in a four-meter-long DN15 pipe with an inner diameter of 17.3 mm. This pipe had a volume of 9.1 liters. Approximately 50% of the total volume was filled with static mixers with a channel diameter of 2 mm. The mixing elements consist of corrugated lamellae forming open, intersecting channels. The channels are arranged in the pipe at intervals of one mixing length, offset by 90°. A continuous circulation of a catalyst phase at a rate of 80 l / h was established through this pipe using a pump. A 2.1% aqueous sodium hydroxide solution was used as the catalyst phase.

[0112] The starting material contained 93.7 wt% n-pentanal and 5.2 wt% 2-methylbutanal; the remaining components, which were 100% absent, consisted of minor components. The starting material was continuously fed into the catalyst cycle at a rate of 8 l / h shortly before the start of the static mixers.

[0113] The two-phase mixture produced at the reactor outlet was separated into an organic product phase and a catalyst phase in a phase separation vessel.

[0114] The system was operated under a nitrogen atmosphere at a pressure of 4 bar, and after three hours of steady-state operation, the following results were obtained. Reaction conditions Temperature °C 130 reactant l / h 2,2 Product composition n-Pentanal (mass %) 3.7 2-Methylbutanal (mass %) 2.6 2-Propylheptenal (mass %) 87.6 4-Propyl-4-Methylhexenal (mass%) 4.6 Remainder (including dissolved water) 1.5 Conversion rate n-pentanal (Xi1) 0.96 Degree of conversion of 2-methylbutanal(Xi2) 0.54 n-pentanal fraction based on the C5 aldehydes in the product (Sn1) 58,2% 2-Propylheptenal fraction based on the unsaturated C10 aldehydes in the product (Sn2) 95.1% Output organic phase including dissolved water kg / h 1.6 Mass flow rate of C5 aldehydes in the output kg / h 0.1 Mass flow rate of α,β-unsaturated C10 aldehydes in the output kg / h 1.5 Example 2

[0115] The starting material used in Example 1, comprising 93.7 wt% n-pentanal and 5.2 wt% 2-methylbutanal and the corresponding minor components, was mixed with the aldolization product from Example 1 in different mass ratios (mixture of partial stream (1b) and aldolization product to starting material stream (2)) and subsequently subjected to hydrogenation.

[0116] For hydrogenation, a recirculating system analogous to DE102009045718A1 Example 1 was used. Similarly, 105 g of catalyst with Cu (6 wt%) / Ni (3.1 wt%) / Cr (0.6 wt%) on Al₂O₃ was used as strand extrudates with a diameter of 1.5 mm and a length of 3–5 mm. The reaction conditions were 180 °C and 25 bar absolute. The reactant feed was 100 ml / h with a recirculation rate of 40 l / h and an exhaust gas rate of 60 l / h. 15% C5 aldehyde / 85% aldolization product 3% C5 aldehyde / 97% aldolization product composition Feed product Feed product n-Pentanal (mass %) 17,2 0,2 6,4 0,1 2-Methylbutanal (mass %) 3,0 0,0* 2,7 0,0* 2-Propylheptenal (mass %) 75,2 0,8 85,9 0,9 4-Propyl-4-Methylhexanal (mass%) 3,9 0,0* 4,5 0,0* n-Pentanol (mass %) 0,0* 16,9 0,0* 6,3 2-Methylbutanol (mass %) 0,0* 3,0 0,0* 2,7 2-Propylheptanol (mass %) 0,0* 73,7 0,0* 84,1 4-Propyl-4-Methylhexanol (mass%) 0,0* 3,8 0,0* 4,4 rest 0,6 1,6 0,5 1,5 n-pentanol content relative to the C5 alcohols in the product 84,9 70,0 2-Propylheptanol content based on the C10 alcohols in the product 95,0 95,0 *below detection limit Example 3

[0117] Simulation studies were conducted to examine the purification of the product stream from hydrogenation. In these studies, a product stream consisting of C10 and C5 alcohols is separated into the individual fractions according to... Figure 2The products were separated using Aspen Plus V7.3 with the PSRK (preedited Soave-Redlich-Kwong) method. The columns are calculated as Radfrac models and designed to achieve product separation with a purity of approximately 99.9%. Four columns with different specifications are used to achieve the desired product qualities. Theoretical separation stages Column I 65 Column II 20 Column III 10 Column IV 55

[0118] The addition of water (in the feed of column I) facilitates the separation of the lighter-boiling components. To prevent water from accumulating in the columns and forming an aqueous phase, water is separated at the top of each column in a decanter. Feed Product stream 1 (alcohol from α,β-aldehydes) Product stream 2 (alcohol from feed aldehydes) 2-Propylheptenal (mass %) 0,9 0 0 n-Pentanol (mass %) 2,8 0 55,5 2-Methylbutanol (mass %) 2,8 0 44,4 2-Propylheptanol (mass %) 87,5 95,4 0 2-Propyl-4-Methylhexanol (mass%) 4,2 4,6 0 H₂O 1,7 0 0 Example 4

[0119] As an alternative to separation using 4 columns, the separation can be carried out using dividing wall columns (also simulated with Aspen Plus V7.3, PSRK and Radfrac). If, for example, columns 2 and 3 are replaced by a 31-stage dividing wall column, the alcohol of the feed aldehydes can be withdrawn as the middle fraction.

[0120] The partition wall extends over 14 of the 31 separation stages. The product is discharged at the 15th stage. The vapor stream striking the partition wall from below is fed equally to the two column segments, while 75% of the liquid stream is directed to the product side and 25% to the feed side.

[0121] If the overhead product of column 1 is directed to the partition column, the following product distribution results: Feed to the partition column Product (side stream of the partition column) H₂O (mass %) 7,4 0 n-Pentanol (mass %) 39,3 66,9 2-Methylbutanol (mass %) 39,5 33,1 2-Propylheptanol (mass %) 0 0 2-Propyl-4-Methylhexanol (mass%) 0,6 0 2-Propylheptenal (mass %) 13,2 0

[0122] The results show that it is possible to control the amounts of C n alcohol and C 2n alcohol using the described procedure, and that the isomer distribution C n can also be controlled in a corresponding way. Reference symbol list

[0123] 1 Hydroformylation product 1a Partial stream of hydroformylation product for aldolization 1b Partial stream of hydroformylation product for hydrogenation 2 Mixture of aldolization product and hydroformylation product 3 Hydrogenation product (with n = 5 essentially water, butane, butene, C5 aldehydes, 2-methylbutanol, pentanol, nonane, α,β-unsaturated C10 aldehydes, C10 aldehydes, α,β-unsaturated C10 alcohols, C10 alcohols, C10+ oligomers with more than 10 carbon atoms) 4 Cn alcohol (with n = 5 essentially 2-methylbutanol, pentanol) 5 C2n alcohol (with n = 5 essentially C10 alcohols) 6. Light current C10 alcohol depleted with proportions of heterogeneous water (e.g., with n=5 essentially water (heterogeneous), butane, butene, C5 aldehydes, 2-methylbutanol, pentanol, nonane, α,β-unsaturated C10 aldehydes, C10 aldehydes, α,β-unsaturated C10 alcohols) 7. Aqueous phase 8. Organic phase C10 alcohol depleted with homogeneous water (with n=5 essentially water (homogeneous),Butane, butene, C5 aldehydes, 2-methylbutanol, pentanol, nonane, α,β-unsaturated C10 aldehydes, C10 aldehydes, α,β-unsaturated C10 alcohols) 9. Light stream (with n=5 essentially water (homogeneous), butane, butene, C5 aldehydes, 2-methylbutanol, pentanol, nonane) 10. Organic phase with homogeneous water (with n=5 essentially water (homogeneous), butane, butene, C5 aldehydes, 2-methylbutanol, pentanol, nonane) 11. Aqueous phase 12. High-boiling stream (with n=5 essentially 2-methylbutanol, pentanol, α,β-unsaturated C10 Aldehydes, C10 aldehydes, α,β-unsaturated C10 alcohols) 14 High-boiling stream (with n= 5 essentially α,β-unsaturated C10 aldehydes, C10 aldehydes, α,β-unsaturated C10 alcohols) 15 High-boiling stream (with n= 5 essentially C10 alcohols, C10+ oligomers with more than 10 carbon atoms) 18 High-boiling stream (with n= 5 essentially C10+ oligomers with more than 10 carbon atoms) 19 Reflux,

Claims

1. Method for preparing saturated Cn- and C2n-alcohols where n = 5 comprising the method steps of a) providing a mixture of isomeric Cn-aldehydes where n = 5, wherein the proportion of unbranched aldehydes is at least 40% by weight, based on the Cn-aldehydes where n = 5, and dividing the mixture into a first substream and a second substream via a split regulator, b) carrying out an aldol condensation of the aldehydes present in the first substream to obtain a mixture of Cn- and α,β-unsaturated C2n-aldehydes, c) mixing the mixture of Cn- and α,β-unsaturated C2n-aldehydes obtained in b) with the second substream provided in a) d) hydrogenating the mixture of Cn- and α,β-unsaturated C2n-aldehydes and isomeric Cn-aldehydes where n = 5 obtained in c) with hydrogen to obtain a mixture of saturated Cn- and C2n-alcohols and e) separating the mixture of Cn and C2n-alcohols via at least one two-column system or via at least one dividing wall column, characterized in that the isomer ratio of the obtained Cn-alcohol is adjusted by the ratio of the first and second substream produced in step a) and by the aldol condensation carried out in step b).

2. Method according to Claim 1, characterized in that the Cn-aldehydes where n = 5 used in step a) were prepared by hydroformylation of isomeric olefins having 4 carbon atoms with synthesis gas in the presence of a hydroformylation catalyst to the corresponding to form the aldehydes specified.

3. Method according to Claim 2, characterized in that in the hydroformylation a catalyst system is used which comprises rhodium as central atom and is complexed with the ligand (1):

4. Method according to one or more of Claims 1 to 3, characterized in that the aldol condensation in step b) is carried out in the presence of aqueous sodium hydroxide solution.

5. Method according to one or more of Claims 1 to 4, characterized in that the aldol condensation of the Cn-aldehydes according to step b) is carried out in a tubular reactor comprising at least one mixing module, which disperses the reactant aldehyde into droplets having an average diameter (Sauter diameter) of 0.2 mm to 2 mm in the continuous catalyst phase (process liquor), which consists of aqueous sodium hydroxide solution and sodium salts of carboxylic acids and has a sodium content of 0.6 to 1.75% by mass and a pH in the range from 12.5 to 13.5.

6. Method according to one or more of Claims 1 to 5, characterized in that the aldol condensation of the Cn-aldehydes according to step b) is carried out in the temperature range from 100 to 150°C.

7. Method according to one or more of Claims 1 to 6, characterized in that the reaction pressure in the reactor during the aldol condensation of the Cn-aldehydes according to step b) is in the range from 0.2 to 1.0 MPa.

8. Method according to one or more of Claims 1 to 7, characterized in that the hydrogenation according to step d) is carried out in the temperature range from 170°C to 200°C at a pressure of 15*105 Pa to 30*105 Pa over a supported catalyst which contains at least nickel and copper as active components.

9. Method according to one or more of Claims 1 to 8, characterized in that the hydrogenation according to step d) is carried out with hydrogen in a pressure range from 5 to 100 bar and the hydrogenation temperatures are between 120 and 220°C.