Continuous method for producing amines in the gas phase using a recycled gas mode

Abstract

1. A process for the continuous production of amines, comprising reacting a primary or secondary alcohol with ammonia in the presence of hydrogen and a heterogeneous hydrogenation catalyst in the gas phase using a recycle gas mode, wherein the temperature in the pressure separator is greater than 20°C.

Description

Technical Field

[0001] The present invention relates to a process for the continuous production of amines, which comprises reacting a primary or secondary alcohol with ammonia in the gas phase in the presence of hydrogen and a heterogeneous hydrogenation catalyst using a recycle gas mode. Background Art

[0002] The products of this process are used in particular as intermediates in the production of fuel additives (US Patent Application Publication No. A-3,275,554; German Patent Application Publication No. A-2125039 and German Patent Application Publication No. A-3611230), surfactants, pharmaceuticals and crop protection agents, curing agents for epoxy resins, catalysts for polyurethanes, intermediates for the preparation of quaternary ammonium compounds, plasticizers, corrosion inhibitors, synthetic resins, ion exchangers, fiber auxiliaries, dyes, vulcanization accelerators and / or emulsifiers.

[0003] WO 2010 / 031719A1 relates to a process for preparing amines in the gas phase. The experimental section teaches the preparation of various amines using the gas recycle mode. According to the examples, the reactor effluent is cooled to 10° C. and fed to a pressure separator (page 22, line 29).

[0004] German Patent Application Publication No. 102004023529A1 (BASF) relates to a process for preparing amines in the gas phase using a heterogeneous catalyst comprising CuO, NiO and Al₂O₃.

[0005] European Patent Application Publication No. 70397A1 (BASF) relates to a process for producing cyclic amines. According to the examples, the reaction mixture is cooled to 20° C. and fed to a pressure separator (page 6, lines 32 to 34).

[0006] German Patent Application Publication No. 19957672A1 (BASF) relates to the production of pyrrolidine, and describes the formation of bis(pyrrolidino)butane as a by-product and its separation from pyrrolidine. [Overview of the project] [Problems that the invention aims to solve]

[0007] The technical problem that this invention aims to solve is to improve existing methods for producing amines from corresponding primary alcohols and to overcome one or more drawbacks of the prior art. The intention was to find a method that is carried out with high conversion rates, high yields including space-time yields, and selectivity.

[0008] The technical challenge was to find a method for the co-production of pyrrolidine and bis(pyrrolidino)butane with high conversion rates, high yields including space-time yields, and selectivity, including efficient means for separation and isolation in high purity. [Means for solving the problem]

[0009] Surprisingly, the above technical challenges can be solved by a method for the continuous production of amines, which involves using a recycle gas mode to react a primary or secondary alcohol with ammonia in the gas phase in the presence of hydrogen and a heterogeneous hydrogenation catalyst, with the temperature in the pressure separator being higher than 20°C.

[0010] It was surprising to learn that the temperature inside the pressure separator had such a significant effect on the selectivity of the desired amine reaction product. Considering that the temperature inside the pressure separator is taught to be 10°C or 20°C, respectively, in the art (see International Publication No. 2010 / 031719A1, page 22, line 29 and European Patent Application Publication No. 70397A1, page 6, lines 32-34), it was surprising that a temperature inside the pressure separator higher than 20°C led to an increase in amine selectivity. [Brief explanation of the drawing]

[0011] [Figure 1] A typical setup of the method according to the present invention is shown. [Modes for carrying out the invention]

[0012] Preferably, the temperature inside the pressure separator is higher than 21°C or even higher than 25°C. The best results in terms of amine selectivity can be obtained by achieving a temperature inside the pressure separator that is higher than 30°C. Preferably, the temperature inside the pressure separator is in the range of 30 to 70°C, more preferably 30 to 60°C.

[0013] This method is carried out in a reactor or multiple reactors. Unless explicitly specified otherwise, the term “reactor” also includes “multiple reactors.” The recycle gas mode is achieved by feeding the reaction mixture obtained in the reactor to a pressure separator, where the reaction mixture is separated into a gas stream and a liquid reaction product stream (also called the “product stream”), and recycling the gas stream back to the reactor. Such a gas stream consists mainly of hydrogen and ammonia (significant amounts of product amines are not found therein). Typically, a portion of such a gas stream is discharged. Otherwise, fresh hydrogen and fresh ammonia would be continuously supplied to the reaction, constantly increasing the amount of gas being processed. The portion of the recycled gas stream is also called “recycled gas” or “recycled gas stream.” The recycled gas stream is typically 40–1500 m 3 Preferably 100-700m 3 (At operating pressure) / [Catalyst m 3 It can have a flow rate in the range of (floor volume)·h.

[0014] Typically, the temperature inside the pressure separator is achieved by cooling the reaction mixture that exits the reactor.

[0015] The pressure separator is typically operated at a pressure close to the reaction pressure, which is further defined below. Typically, the pressure inside the pressure separator is 0.01 to 10 bar (e.g., 0.1 to 10 bar), preferably 0.01 to 5 bar, and particularly preferably 0.5 to 3 bar lower than the reaction pressure.

[0016] The product stream obtained in the pressure separator can be fed into a low-pressure separator, which operates at a pressure lower than the reaction pressure. Typical pressures are in the range of 1 to 10 bar. In the low-pressure separator, the remaining amounts of hydrogen and ammonia, as well as other low-boiling-point substances, corresponding to each amination reaction, are separated. The resulting liquid stream, containing each of the generated amines, can be further separated.

[0017] The product stream obtained in the pressure separator can also be fed into a distillation column, where the remaining amounts of hydrogen and ammonia, as well as other low-boiling substances, corresponding to each amination reaction, can be separated. Preferably, the ammonia is recycled back into the reaction as fresh ammonia. The resulting liquid stream, containing each of the resulting amines, can be further separated. The simultaneous production of pyrrolidine and bis(pyrrolidino)butane will be described in further detail below.

[0018] The reaction takes place in the gas phase. For this purpose, an evaporator can be used, in which each alcohol is evaporated in a gas stream, typically a recycled gas stream. Fresh hydrogen and ammonia can also be supplied directly to the evaporator. Furthermore, fresh hydrogen and / or ammonia can be supplied directly to the recycled gas stream or reactor. A typical setup of the method according to the present invention is shown in Figure 1.

[0019] The method of the present invention is carried out continuously, and the catalyst is preferably installed as a fixed bed in the reactor. Inflow into the fixed catalyst bed is possible from above or below. The temperature, pressure, and volume of the gas flow are set so that relatively high-boiling point reaction products remain in the gas phase.

[0020] Preferably, the reaction is carried out in a tube reactor, in particular a tube bundle reactor or a single stream plant. In the case of a single stream plant, the tube reactor in which the reaction is carried out is preferably composed of a plurality of (for example, two or three) individual tube reactors connected in series. When the reaction is carried out in any such reactor, any reaction pressure or respective range defined herein means the reaction pressure at the reactor inlet. As mentioned above, the pressure in the pressure separator is generally 0.01 to 10 bar (for example, 0.1 to 10 bar), preferably 0.01 to 5 bar, particularly preferably 0.5 to 3 bar lower than the reaction pressure. Accordingly, any respective pressure drop includes the pressure drop occurring over the length of the reactor. A further pressure drop may also occur as a result of cooling of the reaction mixture leaving the reactor before entering the pressure separator.

[0021] Fresh ammonia is added, for example, in a molar amount of 0.90 to 100 times, preferably 1 to 30 times, particularly preferably 1.5 to 10 times, more preferably 2 to 8 times the molar amount of alcohol. It should be understood that these ranges refer to the molar amount of fresh alcohol added to the reaction, and any trace amount of alcohol that may be contained in the recycled gas stream is neglected. For the avoidance of doubt, reference is made to the molar amount of the whole alcohol molecule, not the molar amount of alcohol functional groups. It should be noted that the total amount of ammonia in the reactor results from the amount of fresh ammonia and the amount of ammonia contained in the recycled gas stream that is not considered as fresh ammonia, and therefore exceeds the amount of fresh ammonia. Fresh ammonia may, for example, be any ammonia separated from the product stream in an ammonia column and recycled after leaving the pressure separator.

[0022] Fresh hydrogen is generally added in an amount of 100 to 1000 NL, preferably 150 to 550 NL per (volume L of catalyst and hour), where NL = normal liter, which is the volume converted to STP. STP means standard conditions of temperature and pressure.

[0023] The reaction can be carried out at an absolute pressure in the range of 1 to 300 bar, preferably 10 to 50 bar, particularly preferably 10 to 30 bar, and more preferably 15 to 30 bar.

[0024] The reaction can be carried out at a temperature in the range of 80 to 300°C, preferably 100 to 250°C, particularly preferably 150 to 240°C, and further 170 to 230°C. The reaction can be carried out adiabatically, isothermally or quasi-isothermally (i.e., isoperimetrically), provided that in each case the temperature inside the reactor is within the respective range set forth above. Preferably, the reaction is carried out with an isoperimetric temperature profile, and the reaction temperature is controlled within a range of ±15 K, particularly preferably ±10 K.

[0025] These temperature fluctuations are based on the general temperature in each catalyst bed at the time point when the starting materials enter the catalyst bed and when the reaction mixture exits the catalyst bed, respectively.

[0026] It is also possible to connect a plurality of catalyst beds in parallel or in series.

[0027] When a plurality of catalyst beds are connected in series, the predetermined temperature fluctuation in the isothermal or isoperimetric operation mode according to the present invention applies to the respective temperature inside the catalyst bed at the time point when the starting material enters the first catalyst bed and when the reaction mixture exits the last catalyst bed.

[0028] In a preferred embodiment, as described above, the temperature of the reactor, which is preferably a tubular reactor, is controlled externally by a flow of a heat transfer medium, which may for example be oil, a molten salt or other liquid capable of transferring heat.

[0029] Compared with synthesis in a liquid phase and compared with non-isothermal or isoperimetric synthesis in a gas phase, the reaction conditions according to the present invention have the advantages of particularly good yield and high safety regarding runaway reactions, especially when the reaction temperature is high (for example, 200 to 300°C).

[0030] Isothermal or isocircular gas-phase operation modes significantly reduce the possibility of runaway reactions during synthesis. The mass of material present in the reactor that can be used for runaway reactions is only a small fraction of the mass present in the liquid-phase method.

[0031] The conversion rate of the alcohol is preferably in the range of 80-100%, more preferably 99-100%, and even more preferably 99.5-100%. The conversion rate refers to the molar amount of alcohol consumed in the reaction.

[0032] The liquid space velocity is preferably in the range of 0.1 to 2.0 kg, preferably 0.1 to 1.0 kg, and particularly preferably 0.2 to 0.6 kg of alcohol per liter (bed volume) of catalyst and per hour.

[0033] Since the water produced during the reaction generally does not adversely affect the conversion rate, reaction rate, selectivity, and catalyst operating life, it is advantageous to remove it from the reaction product only during the latter work-up, for example, by distillation.

[0034] The reaction is carried out in the presence of a heterogeneous hydrogenation catalyst. Preferably, the catalytically active composition of such a catalyst before reduction by hydrogen comprises a copper oxygen-containing compound. More preferably, the catalytically active composition of such a catalyst before reduction by hydrogen comprises a copper oxygen-containing compound and another oxidizing agent which is aluminum oxide, zirconium oxide, titanium oxide and / or silicon dioxide. Preferably, the oxidizing agent is aluminum oxide and / or zirconium oxide, and more preferably aluminum oxide.

[0035] To avoid any ambiguity, “and / or” means that each oxidizing agent is any of the listed oxides or a mixture of two or more of the listed oxides, where applicable.

[0036] In a preferred embodiment, the catalytically active composition of the heterogeneous hydrogenation catalyst is used after its final heat treatment and before its reduction with hydrogen. 20-85% by weight of aluminum oxide (Al2O3), zirconium dioxide (ZrO2), titanium dioxide (TiO2), and / or silicon dioxide (SiO2); 1-70% by weight of copper oxygen-containing compounds calculated as CuO; and 0-30% by weight, preferably 0-25% by weight, particularly preferably 0-20% by weight, and especially 0-15% by weight, of nickel oxygen-containing compounds calculated as NiO. Includes.

[0037] In this specification, "and / or" means that each catalyst may contain any of the respective oxides or any mixture thereof.

[0038] The catalytically active composition in such preferred embodiments may further contain 0 to 50% by weight of an oxygen-containing magnesium compound calculated as MgO, an oxygen-containing chromium compound calculated as Cr2O3, an oxygen-containing zinc compound calculated as ZnO, an oxygen-containing barium compound calculated as BaO, and / or an oxygen-containing calcium compound calculated as CaO.

[0039] In the method of the present invention, the catalyst is preferably used in the form of a catalyst that is complete from a catalytically active composition and, optionally, a molding aid (e.g., graphite or stearic acid) when the catalyst is used as a molded body, i.e., it does not contain any further catalytically active additives.

[0040] In this specification, oxidizing substances such as titanium dioxide (TiO2), aluminum oxide (Al2O3), zirconium dioxide (ZrO2), and silicon dioxide (SiO2) are considered to be part of the catalytically active composition.

[0041] The catalyst according to the present invention may, for example, contain an oxygen-containing nickel compound, or may not contain one in nature. To avoid doubt, this also applies to the more preferred, particularly preferred, and very particularly preferred embodiments specified below.

[0042] If an oxygen-containing nickel compound, calculated as NiO, is present, its amount may be, for example, 5 to 28% by weight, preferably 6 to 20% by weight, and particularly preferably 7 to 15% by weight.

[0043] It is also possible to use catalysts that are essentially free of nickel oxygen-containing compounds. In that case, the amount of nickel oxygen-containing compounds calculated as NiO is usually less than 5% by weight, preferably less than 1% by weight, and particularly preferably less than 0.5% by weight (e.g., less than 0.1% by weight). Examples of such catalysts are described in paragraphs

[0101] and

[0102] of German Patent Application Publication No. 102004023529A1.

[0044] To use the catalyst, the catalytically active composition is either introduced into the reaction vessel after being crushed into a powder, or the catalytically active composition is crushed, mixed with a molding aid, molded, and heat-treated, and then placed in the reactor as a molded catalyst, such as pellets, spheres, rings, or extruded products (e.g., extrusion rods).

[0045] Unless otherwise specified, the numerical values ​​(by weight) of the catalyst component concentrations are based on the catalytically active composition of the finished catalyst after the final heat treatment and before reduction by hydrogen, in each respective case.

[0046] The catalytically active composition of the catalyst after its final heat treatment and before it is reduced by hydrogen is defined as the sum of the catalytically active components and the catalyst support material. In the preferred embodiment described above, the catalytically active composition essentially consists of the following components: titanium dioxide (TiO2), and / or aluminum oxide (Al2O3), and / or zirconium dioxide (ZrO2), and / or silicon dioxide (SiO2), and an oxygen-containing compound of copper, and optionally an oxygen-containing compound of magnesium, and / or an oxygen-containing compound of chromium, and / or an oxygen-containing compound of zinc, and / or an oxygen-containing compound of barium, and / or an oxygen-containing compound of calcium, and optionally an oxygen-containing compound of nickel, wherein the amount of these nickel oxygen-containing compounds, calculated as NiO, is 30% by weight or less.

[0047] The total amount of the above components of the catalytically active composition, calculated as Al2O3, ZrO2, TiO2, SiO2, CuO, MgO, Cr2O3, ZnO, BaO, CaO, and NiO, is usually 70-100% by weight, preferably 80-100% by weight, particularly preferably 90-100% by weight, and very particularly preferably 100% by weight.

[0048] The catalytically active composition of the catalyst used in the method of the present invention may further contain one or more elements (oxidation state 0) selected from Groups IA-VIA, B, and VIII of the periodic table, or inorganic or organic compounds thereof.

[0049] Examples of such elements and their compounds include transition metals such as Co and CoO, Re and rhenium oxide, Mn and MnO2, Mo and molybdenum oxide, W and tungsten oxide, Ta and tantalum oxide, Nb and niobium oxide or niobium oxalate, V and vanadium oxide, vanadyl pyrophosphate; lanthanides such as Ce and CeO2 or Pr and Pr2O3; alkali metal oxides such as Na2O; alkali metal carbonates; alkaline earth metal oxides such as SrO; alkaline earth metal carbonates such as MgCO3, CaCO3 and BaCO3; and boron oxide (B2O3).

[0050] In a more preferred embodiment, the catalytically active composition of the catalyst used in the method of the present invention is, after its final heat treatment and before its reduction with hydrogen, 25-80% by weight, preferably 30-75% by weight, of aluminum oxide (Al2O3), zirconium dioxide (ZrO2), titanium dioxide (TiO2), and / or silicon dioxide (SiO2); 2 to 65% by weight, preferably 5 to 60% by weight, particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO; and 0-30% by weight, preferably 0-25% by weight, particularly preferably 0-20% by weight, and especially 0-15% by weight, of nickel oxygen-containing compounds calculated as NiO. Includes.

[0051] In a particularly preferred embodiment, the catalytically active composition of the catalyst used in the method of the present invention is, after its final heat treatment and before its reduction with hydrogen, 25-80% by weight, preferably 30-75% by weight, of aluminum oxide (Al2O3) and / or zirconium dioxide (ZrO2); 2 to 65% by weight, preferably 5 to 60% by weight, particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO; and 0-30% by weight, preferably 0-25% by weight, particularly preferably 0-20% by weight, and especially 0-15% by weight, of nickel oxygen-containing compounds calculated as NiO. Includes.

[0052] In these particularly preferred embodiments, the presence of titanium dioxide (TiO2) and silicon dioxide (SiO2) is not excluded. When the catalyst also contains titanium dioxide (TiO2) and / or silicon dioxide (SiO2), the total amounts of aluminum oxide (Al2O3), zirconium dioxide (ZrO2), titanium dioxide (TiO2), and / or silicon dioxide (SiO2) are preferably 25-80% by weight or 30-75% by weight, respectively.

[0053] In a very particular embodiment, the catalytically active composition of the catalyst used in the method of the present invention is, after its final heat treatment and before its reduction with hydrogen, 25-80% by weight, preferably 30-75% by weight, of aluminum oxide (Al2O3); 2 to 65% by weight, preferably 5 to 60% by weight, particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO; and 0-30% by weight, preferably 0-25% by weight, for example 0-20% by weight, particularly 0-15% by weight, of an oxygen-containing compound with nickel as NiO. Includes.

[0054] In such a particularly preferred embodiment, the presence of zirconium dioxide (ZrO2), titanium dioxide (TiO2), and silicon dioxide (SiO2) is not excluded. If the catalyst also contains zirconium dioxide (ZrO2), titanium dioxide (TiO2), and / or silicon dioxide (SiO2), the total amounts of aluminum oxide (Al2O3), zirconium dioxide (ZrO2), titanium dioxide (TiO2), and / or silicon dioxide (SiO2) are preferably 25-80% by weight or 30-75% by weight, respectively.

[0055] In the more preferred, particularly preferred, and very particularly preferred embodiments described above, the catalytically active composition of the catalyst used in the method of the present invention may further contain 0 to 30% by weight, preferably 0 to 20% by weight, of an oxygen-containing magnesium compound calculated as MgO, and / or an oxygen-containing chromium compound calculated as Cr2O3, and / or an oxygen-containing zinc compound calculated as ZnO, and / or an oxygen-containing barium compound calculated as BaO, and / or an oxygen-containing calcium compound calculated as CaO.

[0056] The oxygen-containing copper compounds are particularly copper(I) oxide and copper(II) oxide, preferably copper(II) oxide.

[0057] In a very preferred embodiment, the catalytically active composition of the catalyst is after its final heat treatment and before its reduction with hydrogen. 25-80% by weight, preferably 30-75% by weight, of aluminum oxide (Al2O3); and 2 to 65% by weight, preferably 5 to 60% by weight, and particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO. It becomes essentially.

[0058] In another highly preferred embodiment, the catalytically active composition of the catalyst is after its final heat treatment and before its reduction with hydrogen. 25-80% by weight, preferably 30-75% by weight, aluminum oxide (Al2O3), 2 to 65% by weight, preferably 5 to 60% by weight, particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO; and A nickel oxygen-containing compound calculated as NiO, less than 30% by weight, preferably 5 to 28% by weight, more preferably 6 to less than 20% by weight, and even more preferably 7 to less than 15% by weight. It becomes essentially.

[0059] "Essentially consisting of" means that the catalytic active composition of the catalyst consists of more than 95% by weight, preferably more than 99% by weight, of aluminum oxide, copper oxygen-containing compounds, and optionally nickel oxygen-containing compounds.

[0060] The catalyst used in the method of the present invention can be prepared by various methods. For such methods, see German Patent Application Publication No. 102004023529A1, incorporated herein by reference, particularly paragraphs

[0046] to

[0063] .

[0061] The method according to the present invention is particularly suitable for the production of each cyclic amine, and the method consists of the reaction of a diol with ammonia.

[0062] Preferred diols are primary aliphatic diols having 2 to 6, preferably 4 to 6, carbon atoms. In this specification, the term “aliphatic” means non-cyclic, non-aromatic functionalized or non-functionalized organic residues that do not contain aromatic ring systems. These may have any heteroatoms, such as oxygen, as part of any functional group.

[0063] Preferred aliphatic diols are selected from the group consisting of diethylene glycol, 1,4-butanediol, 1,5-pentanediol, and 1,6-hexanediol. 1,4-butanediol is particularly preferred.

[0064] In such reactions between diols and ammonia, even more valuable products can be produced depending on the specific reaction conditions.

[0065] The method according to the present invention is also particularly suitable for the production of methoxy-2-propylamine, which comprises reacting 1-methoxy-2-propanol with ammonia.

[0066] Furthermore, the method according to the present invention is particularly suitable for the production of mono-, di-, and trihexylamines, the method comprising reacting hexane-1-ol with ammonia.

[0067] The method according to the present invention is particularly suitable for the simultaneous production of pyrrolidine and bis(pyrrolidino)butane, and the method comprises reacting 1,4-butanediol with ammonia.

[0068] Without wishing to be bound by any theory or without limiting the scope of the present invention in any way, the simultaneous production of pyrrolidine and bis(pyrrolidino)butane is thought to occur via the following reaction scheme. [ka]

[0069] The ratio of pyrrolidine to bis(pyrrolidino)butane can be varied depending on the reaction conditions and method setup. Operating the pressure separator within the temperature range according to the present invention improves the yield of pyrrolidine, which is more economically favorable.

[0070] Alternatively, or in addition to recycling pyrrolidine, fresh pyrrolidine can be supplied to the reactor along with ammonia and alcohol.

[0071] In the case of the simultaneous production of pyrrolidine and bis(pyrrolidinyl)butane, the product stream obtained from the pressure separator contains, in particular, pyrrolidine, bis(pyrrolidinyl)butane, 4-hydroxybutylpyrrolidine, 4-aminobutylpyrrolidine, a high-boiling-point substance having a boiling point higher than that of bis(pyrrolidinyl)butane, ammonia, and water. Preferably, such a product stream is further separated, for example, by distillation. It is in principle possible to recycle the pyrrolidine thus removed from the product stream into a reaction that leads to increased bis(pyrrolidinyl)butane production. Without wishing to be bound by any theory or without limiting the scope of the present invention in any way, such recycling is considered to increase the concentration of pyrrolidine in the reactor, as can be seen from the reaction scheme described above, and thereby favor the reaction to bis(pyrrolidinyl)butane.

[0072] Figure 1 shows a particularly preferred embodiment. Alcohol is supplied to the evaporator (4) via line (1) and recycled gas via line (2). The recycled gas is passed through the compressor (11) to raise the pressure to the desired reaction pressure. Fresh hydrogen and ammonia can be supplied directly into the evaporator (4) via lines (3a) and (3a') or into the recycled gas via lines (3b) and (3b'). They can also be supplied into the recycled gas via lines (3c) and (3c') before passing through the compressor (11). This is advantageous because the amount of energy required for the compressor can be reduced because both the hydrogen and ammonia flows have a higher pressure than the recycled gas. Hydrogen and ammonia can also be supplied directly into the reactor via lines (3d) and (3d'). Theoretically, any combination of such methods of adding hydrogen (i.e., (3a)~(3d)) and ammonia (i.e., (3a')~(3d')) is also possible. In the evaporator (4), the alcohol is evaporated, and the resulting gas stream is supplied to the reactor (6) via line (5). The reaction mixture is passed through a heat exchanger (8), and optionally through line (7) to a cryostat (not shown in Figure 1), where it is cooled and supplied to a pressure separator (9), where a gas stream essentially consisting of hydrogen and ammonia is extracted. A portion of the gas stream is discharged via line (10). The remainder is recycled as recycled gas back to the evaporator (4) via line (2).

[0073] Furthermore, the crude reaction products can be supplied from the pressure separator (9) to the low-pressure separator (13) via line (12) for further degassing. The resulting gas stream consists essentially of hydrogen, ammonia, and their respective low-boiling substances and is discharged via line (15). The crude reaction products, particularly the valuable generated amines and high-boiling substances, are removed from the low-pressure separator (13) via line (14). The crude amine products can be further purified.

[0074] The crude reaction product can also be supplied from the pressure separator (9) to a distillation column (not shown in Figure 1) via line (12), where hydrogen, ammonia, and their respective low-boiling substances are removed. Preferably, the ammonia is recycled back into the reaction as fresh ammonia.

[0075] Typically, low-pressure separators are used in laboratory settings, while distillation columns are used on an industrial scale.

[0076] The following is one embodiment of the present invention. (1) A method for the continuous production of amines, comprising using a recycle gas mode to react a primary or secondary alcohol with ammonia in the gas phase in the presence of hydrogen and a heterogeneous hydrogenation catalyst, wherein the temperature in the pressure separator is higher than 20°C. (2) The method according to (1), wherein the temperature inside the pressure separator is higher than 21°C, preferably higher than 25°C, particularly preferably higher than 30°C, and more preferably in the range of 30 to 70°C or 30 to 60°C. (3) The method according to (1) or (2), wherein the pressure separator is operated at a pressure close to the reaction pressure. (4) The method according to (3), wherein the pressure in the pressure separator is 0.01 to 10 bar (for example, 0.1 to 10 bar), preferably 0.01 to 5 bar, and particularly preferably 0.5 to 3 bar lower than the reaction pressure. (5) The recycled gas flow is 40-1500 m (at operating pressure). 3 Preferably 100-700m 3 / [catalyst m 3 The method according to any one of (1) to (4), having a flow rate in the range of (floor volume)·h. (6) The method according to any one of (1) to (5), wherein fresh ammonia is added in a molar amount of 0.90 to 100, preferably 1 to 30, particularly preferably 1.5 to 10, or even 2 to 8 times the molar amount of the alcohol. (7) The method according to any one of (1) to (6), wherein the reaction is carried out at an absolute pressure in the range of 1 to 300 bar, preferably 10 to 50 bar, particularly preferably 10 to 30 bar, or even more preferably 15 to 30 bar. (8) The method according to any one of (1) to (7), wherein the reaction is carried out at a temperature in the range of 80 to 300°C, preferably 100 to 250°C, particularly preferably 150 to 240°C, or even more preferably 170 to 230°C. (9) The method according to any one of (1) to (8), wherein the liquid space velocity is in the range of 0.1 to 2.0 kg, preferably 0.1 to 1.0 kg, and particularly preferably 0.2 to 0.6 kg of alcohol per liter (bed volume) of catalyst and per hour. (10) The catalytically active composition of the heterogeneous hydrogenation catalyst is after its final heat treatment and before its reduction with hydrogen. 20-85% by weight of aluminum oxide (Al2O3), zirconium dioxide (ZrO2), titanium dioxide (TiO2), and / or silicon dioxide (SiO2); 1-70% by weight of copper oxygen-containing compounds calculated as CuO; and 0-30% by weight, 0-25% by weight, and particularly preferably 0-20% by weight, of nickel oxygen-containing compounds calculated as NiO. The method described in any of (1) to (9), including the method described in (1) to (9). (11) The catalytically active composition of the heterogeneous hydrogenation catalyst is after its final heat treatment and before its reduction with hydrogen. 25-80% by weight, preferably 30-75% by weight, of aluminum oxide (Al 2 O 3 ) and / or zirconium dioxide (ZrO 2 ); 2 to 65% by weight, preferably 5 to 60% by weight, particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO; and 0-30% by weight, preferably 0-25% by weight, and particularly preferably 0-20% by weight, of an oxygen-containing nickel compound calculated as NiO. The method described in any of (1) to (9), including the method described in (1) to (9). (12) The catalytically active composition of the heterogeneous hydrogenation catalyst is after its final heat treatment and before its reduction with hydrogen. 25-80% by weight, preferably 30-75% by weight, of aluminum oxide (Al 2 O 3 ); 2 to 65% by weight, preferably 5 to 60% by weight, particularly preferably 20 to 60% by weight, of an oxygen-containing copper compound calculated as CuO; and 0-30% by weight, preferably 0-25% by weight, and particularly preferably 0-20% by weight, of an oxygen-containing nickel compound calculated as NiO. The method described in any of (1) to (9), including the method described in (1) to (9). (13) The method according to any one of (10) to (12), wherein the catalytic active composition of the heterogeneous hydrogenation catalyst contains 5 to 28% by weight, preferably 6 to 20% by weight, and particularly preferably 7 to 15% by weight, of an oxygen-containing nickel compound calculated as NiO, after the final heat treatment and before its reduction by hydrogen. (14)- Each cyclic amine, wherein the method comprises reacting a diol, preferably an aliphatic primary diol having 2 to 6 carbon atoms, particularly preferably diethylene glycol, 1,4-butanediol, 1,5-pentanediol or 1,6-hexanediol with ammonia, - 1-methoxy-2-propylamine, wherein the method comprises reacting 1-methoxy-2-propanol with ammonia, and - Mono-, di- and trihexylamines, wherein the method comprises reacting hexane-1-ol with ammonia. The method described in any of (1) to (13) for the manufacture of [the product]. (15) The method according to any one of (1) to (14) for the simultaneous production of pyrrolidine and bis(pyrrolidino)butane, comprising reacting 1,4-butanediol with ammonia. The following examples are for the purposes described in the present invention only and do not limit the present invention in any way. [Examples]

[0077] catalyst: The following examples were carried out using a copper / nickel catalyst having a composition of 45 wt% CuO and 10 wt% NiO, with the remainder up to 100% being gamma-Al2O3 (after its final heat treatment and before its reduction with hydrogen).

[0078] The catalyst was prepared according to Example 1 of German Patent Application Publication No. A-2445303. The catalyst was reduced before the reaction commenced (see below).

[0079] Examples 1-4 The experiment was carried out continuously in a gas-phase reactor. In this reactor, the reactants flowed from bottom to top through an oil-heated double-walled tube with an inner diameter of 4.8 cm and a length of 2.1 m. 40 ml of ceramic spheres (2.5-3.5 mm), 1 liter of catalyst, and 1.5 liters of inert material (ceramic spheres, 2.5-3.5 mm) were packed from bottom to top. The reactor was operated at 20 bar. The pelletized catalyst was used in a size of 5 x 5 mm (i.e., 5 mm in diameter and 5 mm in height). After installation in the reactor, all catalysts were activated at atmospheric pressure according to the following method: 12 hours at 180°C (oil circulation reactor) with 20 NL / H of N2 and 400 NL / H of N2; 12 hours at 200°C with 20 NL / H of 40H2 and 400 NL / H of N2; 6 hours of N2 replacement with 200 NL / H of H2; and 6 hours at 240°C with 200 NL / H of H2. (NL = standard liters = volume converted to STP). The feed streams of fresh hydrogen, circulating gas, pressurized gas, and starting materials were heated to the desired reactor temperature by a system including three coil heat exchangers. The third heat exchanger was controlled by a temperature sensor just before the reactor. The oil heating in the double-wall reactor was similarly set to the desired reactor temperature. Furthermore, through two coil-type heat exchangers, the reactor products were first cooled by river water, then heated to the desired temperature (25–49°C) in the pressure separator using a cryostat, and supplied to the pressure separator. The pressure separator was operated at a pressure approximately 2.5 bar lower than the reaction pressure (20 bar). There, separation of the liquid phase and gas phase occurred. The liquid phase was depressurized in a low-pressure separator maintained at 45°C, and the gas released from there was discharged via off-gas, while the liquid was sent to the product drum. The gas phase from the pressure separator was recirculated in a predetermined amount via a circulating gas compressor and again acted as a carrier gas for the starting materials. By a pressure regulator, the excess gas was sent to a muffle furnace for incineration. The conversion rate and selectivity of the products were determined by gas chromatography analysis and reported in corrected GC area %.

[0080] The setup described above corresponds to the setup explained in Figure 1.

[0081] The reaction conditions for Examples 1 to 4 are as follows: Reactor inlet temperature: 199℃ Reactor outlet temperature: 209℃ Evaporator temperature: 220℃ Reaction pressure: 20 bar Liquid space velocity: 0.5kg BDO / (L(Cat.) h) Molar ratio (fresh NH3:BDO): 3:1 Hydrogen flow rate: 150 NL / [Catalyst L (floor volume)]·h] Recycled gas flow rate: 7 Nm 3 / [catalyst L(bed volume))] h] Conversion BDO: 100% (Nm 3 (=Standard cubic meter = Volume converted to STP)

[0082] The results are shown in Table 1 below.

[0083] [Table 1]

[0084] Analysis of the results As shown in Table 1, increased pyrrolidine selectivity can be achieved by raising the temperature of the pressure separator. A plateau is reached in the temperature range of 25–35°C. Furthermore, the overall selectivity for amine products (especially the valuable products pyrrolidine and bis(pyrrolidino)butane) increases. In addition, the amount of unwanted by-product THF is consistently reduced.

Claims

1. A method for the continuous production of pyrrolidine and bis(pyrrolidino)butane, comprising: feeding a reaction mixture obtained in a reactor to a pressure separator, where the reaction mixture is separated into a gas stream and a liquid reaction product stream, and using a recycled gas mode achieved by recycling the gas stream back to the reactor, reacting 1,4-butanediol with ammonia in the gas phase in the presence of hydrogen and a heterogeneous hydrogenation catalyst at an absolute pressure in the range of 1 to 50 bar, wherein the temperature in the pressure separator is higher than 25°C, the pressure in the pressure separator is 0.01 to 10 bar lower than the reaction pressure, and fresh ammonia is added in a molar amount of 1.5 to 100 times the molar amount of 1,4-butanediol.

2. The method according to claim 1, wherein the temperature inside the pressure separator is higher than 30°C.

3. The method according to claim 1 or 2, wherein the pressure in the pressure separator is 0.1 to 10 bar lower than the reaction pressure.

4. The recycled gas flow rate is 40-1500 m³ / [catalyst m³] (at operating pressure). ３ The method according to any one of claims 1 to 3, having a flow rate in the range of (floor volume) * h.

5. The method according to any one of claims 1 to 4, wherein fresh ammonia is added in a molar amount 1.5 to 10 or 2 to 8 times the molar amount of 1,4-butanediol.

6. The method according to any one of claims 1 to 5, wherein the reaction is carried out at an absolute pressure in the range of 10 to 30 bar or 15 to 30 bar.

7. The method according to any one of claims 1 to 6, wherein the reaction is carried out at a temperature of 80 to 300°C.

8. The method according to any one of claims 1 to 7, wherein the liquid space velocity is in the range of 0.1 to 2.0 kg of 1,4-butanediol per liter (bed volume) of catalyst and per hour.

9. The catalytically active composition of the heterogeneous hydrogenation catalyst is, after its final heat treatment and before its reduction by hydrogen, 20 to 85% by weight of aluminum oxide (Al₂O₃), zirconium dioxide (ZrO₂), titanium dioxide (TiO₂), and / or silicon dioxide (SiO₂); 1 to 70% by weight of copper oxygen-containing compounds calculated as CuO; and 0-30% by weight of nickel oxygen-containing compounds calculated as NiO The method according to any one of claims 1 to 8, including

10. The catalytically active composition of the heterogeneous hydrogenation catalyst is, after its final heat treatment and before its reduction by hydrogen, 25-80% by weight of aluminum oxide (Al ２ O ３ ) and / or zirconium dioxide (ZrO ２ ); 2 to 65% by weight of copper oxygen-containing compounds calculated as CuO; and 0-30% by weight of nickel oxygen-containing compounds calculated as NiO The method according to any one of claims 1 to 8, including

11. The catalytically active composition of the heterogeneous hydrogenation catalyst is, after its final heat treatment and before its reduction by hydrogen, 25-80% by weight of aluminum oxide (Al ２ O ３ ); 2 to 65% by weight of copper oxygen-containing compounds calculated as CuO; and 0-30% by weight of nickel oxygen-containing compounds calculated as NiO The method according to any one of claims 1 to 8, including

12. The method according to any one of claims 9 to 11, wherein the catalytic active composition of the heterogeneous hydrogenation catalyst comprises 5 to 28% by weight of an oxygen-containing nickel compound calculated as NiO after the final heat treatment and before its reduction by hydrogen.

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