WO2024115263A1

WO2024115263A1 - Method for manufacture of isophoronediamine

Info

Publication number: WO2024115263A1
Application number: PCT/EP2023/082817
Authority: WO
Inventors: Martin SCHLODERER; Alfred Krause; Marina-Eleni STAVROU
Original assignee: Basf Se
Priority date: 2022-12-02
Filing date: 2023-11-23
Publication date: 2024-06-06

Abstract

A process for the manufacture of IPDA by hydrogenation of IPNI and/or IPN, comprising the steps of: a) feeding a stream comprising IPNI and/or IPN to a first hydrogenation reactor or a first series of hydrogenation reactors, b) feeding at least a part of the effluent from the first hydrogenation or the first series of hydrogenation reactors to a second hydrogenation reactor or a second series of hydrogenation reactors, and c) recycling at least a part of the effluent from the first hydrogenation or the first series of hydrogenation reactors to the first hydrogenation reactor or the first series of hydrogenation reactors.

Description

Method for Manufacture of Isophoronediamine

Description

The present invention relates to a process for the manufacture of isophorone diamine (I PDA).

IPDA is used as a starting product for preparing isophorone diisocyanate (IPDI), an isocyanate component for polyurethane systems, as an amine component for polyamides and as a hardener for epoxy resins.

IPDA is usually prepared in a multistage process starting from isophorone (IP). In a first step, hydrogen cyanide (HCN) is added to IP to obtain the corresponding isophorone nitrile (IPN). In a further step, IPN is converted to IPDA by converting the carbonyl group of IPN to an amino group and the nitrile group to an aminomethyl group in the presence of ammonia, hydrogen and hydrogenation catalysts. The second step can be divided into further steps, in which the carbonyl group of IPN is first converted with ammonia (NH3) to the corresponding isophorone nitrile imine (I PN I) in the presence of an imination catalyst. In a subsequent step, IPNI is then hydrogenated in the presence of a hydrogenation catalyst to obtain IPDA.

In an IPDA manufacturing process, it is desirable to achieve the following objectives:

A) Manufacture of IPDA to obtain a high product yield of IPDA.

B) Manufacture of IPDA having a high isomer ratio of cis-lPDA to trans-lPDA.

The control of the cis-trans-ratio (CTR) is important since the different IPDA isomers have different reactivities in down-stream applications, e.g., when used as an educt for the manufacture of epoxides or polyurethanes. According to DE-A-4211454, IPDA with a high cis-trans-ratio (CTR) of 75:25 and more are preferred in applications, which require a short pot-life and a short curing temperature. This is the case in most epoxy and PUR-applications. IPDA with a high CTR is therefore commercially preferred. Some customers specify a CTR of >75:25 for their applications.

The CTR in IPDA is influenced by many factors.

One prior art process discloses that a high CTR can be achieved by a two-stage conversion of IPNI by controlling the temperature in the respective stages (EP 0394968).

According to DE 19507398 and DE19747913, the addition of a base or a basic compound to the hydrogenative amination also has an influence on the isomer ratio.

W02008077852 further teaches that the time of addition of the base to the hydrogenation step can also lead to an increase of the CTR.

High CTRs are also achieved when the hydrogenation reaction is carried out with basic catalysts (DE4010227 and EP0623585).

An increase of the CTR was also reported when the reductive amination was carried out in the presence of an acid (DE19756400).

Even if the reaction conditions are carefully selected to control the CTR, e.g. by the choice of catalyst, it is possible that the CTR decreases when the catalysts employed in the reaction ages and loses at least some of its selectivity towards cis-lPDA.

To counterbalance the decrease of the CTR, it is sometimes proposed in the state of the art to subject the produced IPDA to an isomerization step (WO2016143538, EP1529028).

DE10236674 teaches a method for enhancing the CTR by distillation. The method makes use of the principle that the cis isomer of IPDA has a higher boiling point than the trans-isomer. A crude IPDA having a CTR of less than 73:27 is separated into a fraction having a CTR of <66:34, which may be drawn-off at the top of the distillation column, and a fraction having a CTR of >73:27, which is typically drawn-off as a side-offtake. The distillation parameters, such a reflux and temperature, are controlled to achieve the quality of the desired fractions. The process according to DE10236674 has the advantage, that an IPDA fraction having a high CTR can be obtained, which can be used in applications requiring a high CTR, while further obtaining a fraction with a lower CTR, which can be used in application in which the CTR is of lower importance. Using the process of DE10236674, nearly the entire yield of IPDA produced can be utilized without substantial losses.

However, achieving a high CTR is not the only objective when designing an efficient IPDA manufacturing process.

In addition to achieving a high CTR, it is desirable to obtain IPDA itself in a high yield.

Obtaining high yields of IPDA requires high degrees of conversion of IPNI and a high selectivity of the conversion of IPNI to IPDA to avoid unconverted IPNI and undesirable side products formed during the conversion of IPNI in the product stream.

One kind of undesirable species in the product streams are nitriles.

One kind of nitrile species which can be present in the product stream is unconverted IPN. Another kind of nitrile species which can be present in the product stream is IPNA.

IPNA is obtained by hydrogenation of the imine moiety of IPNI. IPNA remains in the product stream if IPNA if the nitrile group is not further hydrogenated to obtain IPDA.

Nitriles are difficult to separate from the product stream and can itself participate in further reactions resulting in side-products which can impart unwanted product properties, such as odor and discoloration (yellowing) in downstream applications.

IPNA itself can react with other amines to form undesired secondary and tertiary amines.

IPNA can also undergo an intramolecular ring closure to form an amidine intermediate which is further converted to a ring species of formula (1):

In the literature, different approaches have been disclosed for achieving high conversions of IPNI and/or IPN and high yields of IPDA.

WO2012126956 (BASF) discloses the hydrogenation of IPNI and/or IPN in a fixed-bed hydrogenation reactor in which the cross-sectional loading, based on the liquid phase passing through the reactor, is in the range of 5 kg/(m²s) to 50 kg/(m²s). The high cross-sectional loading is achieved by utilizing reactors with a slim design, i.e. , reactors in which the ratio of height to diameter of the reactor is in the range of 1 :1 to 500:1. In a further preferred embodiment, the high cross-sectional loading is achieved by recycling a part of the product stream from the hydrogenation reactor. Preferably, the ratio of the recycling stream to the feed stream is in the range of 0.5:1 to 250:1.

EP3406589 (EVON IK) discloses the hydrogenation of IPNI and/or IPN in a fixed bed hydrogenation reactor in which the cross-sectional loading of the hydrogenation reactor, based on the liquid phase passing through the reactor, is less or equal to 4 kg/(m²s). The low cross-sectional loading is achieved by either a single pass of the feed stream through the reactor, i.e., without a recycling a part of the product stream, or by recycling only a small portion of the product stream. Preferably, the ratio of the recycling stream to the feed stream is in the range of 0:1 to 0.49:1.

A main difference between operating reactors with a different cross-sectional loading is the degree of back-mixing which may occur in the reactors because of the degree of cross-sectional loading in the reactor.

In reactors operated with a high cross-sectional loading, the flow characteristic of the feed stream is expected to trend toward the flow characteristic of a continuously stirred tank reactor (CSTR). In reactors operated with a low cross-sectional loading, the flow characteristic of the feed stream is expected to trend toward the flow characteristic of a plug-flow reactor.

Accordingly, the higher the cross-sectional load of a reactor, the higher the degree of mixing and back-mixing within the reactor.

Back-mixing broadens the residence time distribution of volume elements within a reactor and leads to a more even temperature profile along the axis of the reactor, resulting in a more isothermal profile.

Plug-flow with little or no back-mixing results in a narrow residence time distribution and a temperature profile with a temperature maximum between the inlet and the outlet of the reactor. A fixed bed reactor operated with a high cross-sectional loading and a broader residence time distribution is expected to have a lower degree of conversion of IPNI and partially hydrogenated IPNA, resulting in lower yields of IPDA and higher concentrations of IPNI and IPNA.

Contrary to these expectations it was found in WO2012126956 that increasing the high cross- sectional loading did not result in an increase of undesired IPNA in the product stream. Rather, higher cross-sectional loadings led to a decrease of the IPNA concentration and an increase of the overall IPDA yield.

However, the overall yield of IPDA obtained in a reactor operated with a high cross-sectional loading appears to be less than the yield of IPDA achieved when operating the hydrogenation reactor in a mode of low cross-sectional loadings, as done in EP3406589. Low cross-sectional loading however may result in hotspots by insufficient wetting of the catalyst bed or insufficient mixing leading to a weak convection of heat. These localized hotspots can promote the formation of undesired side products leading to an unwanted decrease of the selectivity for IPDA.

In view of prior art, there remains the need for a process for the productions of IPDA, which fulfills both needs, namely

A) the manufacture of IPDA to obtain a high product yield of IPDA, and

B) the manufacture of IPDA having a high isomer ratio of cis-lPDA to trans-lPDA.

The problem of the present invention was solved by a process for the manufacture of IPDA by hydrogenation of IPNI and/or IPN, comprising the steps of: a) feeding a stream comprising IPNI and/or IPN to a first hydrogenation reactor or a first series of hydrogenation reactors, b) feeding at least a part of the effluent from the first hydrogenation or the first series of hydrogenation reactors to a second hydrogenation reactor or a second series of hydrogenation reactors, and c) recycling at least a part of the effluent from the first hydrogenation or the first series of hydrogenation reactors to the first hydrogenation reactor or the first series of hydrogenation reactors.

In the process according to the present invention, the hydrogenation of IPNI and/or IPN is carried out in two or two series of hydrogenation reactors. The process is characterized therein that part of the effluent from the first hydrogenation reactor or the first series of hydrogenation reactors is recycled and only the other part of the effluent, which is not recycled, is fed to a second or a second series of hydrogenation reactors.

The process is considered to differ from WO2012126956 therein that WO2012126956 discloses a recycling stream from the second to the first hydrogenation reactor. WO2012126956 does not disclose a second hydrogenation reactor or a series of hydrogenation reactors downstream of where the recycle stream from the first reactor or first series of reactors branches off.

Also, EP3406589 discloses that that the hydrogenation is preferably carried out in a reactor where the educts pass through the reactor in a single pass, i.e. without recycling the effluent from the hydrogenation reactor to the feed.

Surprisingly it has been found that a particularly efficient process for the manufacture of IPDA can be obtained if - compared to the teaching of WO2012126956 - the process for the manufacture of IPDA comprises a second hydrogenation reactor or a second series of hydrogenation reactors which are located downstream of where the recycle stream from the first hydrogenation reactor or the first series of hydrogenation reactors to the feed of the first hydrogenation reactor or the first series of hydrogenation reactors branches off.

The process according to the invention has the following advantages: the first hydrogenation reactor or the first series of hydrogenation reactors can be operated with a higher cross-sectional loading. This higher cross-sectional loading results in a more homogeneous temperature profile in the reactors. Hotspot formation may be significantly reduced. The avoidance of hotspots potentially results in fewer side products, increasing the selectivity and yield of IPDA. In addition, lower temperatures favor a high CTR, which is often desired in IPDA production. through the recycle stream, ammonia can be recycled to the reactor. In addition, the ammonia load to the downstream IPDA refining section, where hydrogen, ammonia and side products are separated can be reduced by the recycle stream. Further, the temperature of the recycle stream may be increased when two or two series of hydrogenation reactors are used. It was found that the hydrogen solubility was increased with increasing temperature. A higher hydrogen solubility in the liquid phase generally results in higher conversion rates and reduced side product formation. the second hydrogenation reactor or the second series of hydrogenation reactors can be operated with a lower cross-sectional loading. Thus, the residence time in the reactors can be increased, which generally leads to higher conversions. coupling a first hydrogenation reactor or a first series of hydrogenation reactors with a second or a second series of hydrogenation reactors according to the invention was also surprisingly found to reduce the overall energy duty required for the operation of the thermal equipment. the equipment for cooling and separating the liquid and the gas phase in the downstream process after the last hydrogenation step can be designed with smaller dimensions.

In the process according to the present invention IPNI and/or IPN is hydrogenated to IPDA.

The process according to the invention comprises the step a) of a) feeding a stream comprising IPNI and/or IPN to a first hydrogenation reactor or a first series of hydrogenation reactors.

It is possible to feed the IPN usually obtained from the reaction of hydrogen cyanide (HCN) with IP directly into the first or first series of hydrogenation reactors.

In a preferred embodiment, IPN is fully or partially converted to IPNI by imination before feeding IPNI or a mixture of IPNI and unconverted IPN to the first hydrogenation reactor or a first series of hydrogenation reactors.

Imination:

The imination of IPN is usually conducted at temperature from 20 to 150°C, preferably 30 to 100°C and more preferably 50 to 90°C and a pressure of 50 to 300 bar, preferably 100 to 250 bar and more preferably 150 to 220 bar.

Suitable imination catalysts are usually acidic oxides, preferably alumina, titania, zirconia and silica.

The catalyst loading is preferably in the range of 0.01 to 10, more preferably 0.05 to 7 and even more preferably 0.1 to 5 kg IPN per kg catalyst.

The molar ratio of NH3 to IPN is usually in the range of 5:1 to 500:1 , preferably 10:1 to 400:1 and more preferably 20:1 to 300:1.

The imination can be optionally conducted in the presence of a solvent, such as alcohols or ethers, in particularly THF, ethanol or butanol. Most preferably, the imination is not conducted in the presence of a solvent.

The imination can be conducted in one or more pressurized reaction vessels, most preferably the one or more pressurized reaction vessels are one or more tubular reactors where the imination catalyst is arranged in a fixed bed. Preferably the imination is conducted in 1 to 3, more preferably 1 to 2 and even more preferably in one reactor.

The reaction conditions, such as temperature, catalyst, pressure, reactor geometry, are selected in such a manner that the conversion of I PN to IN PI is preferably 80% or more, more preferably 90% or more and most preferably 95% or more.

Hydrogenation:

The process according to the present invention comprises feeding a stream comprising IPNI and/or IPN to a first hydrogenation reactor or a first series of hydrogenation reactors.

Preferably the feed comprises 80 percent by weight or more of IPNI, more preferably 90 percent by weight and most preferably 95 percent by weight or more. As specified above, streams comprising IPN and IPNI can be obtained by imination of IPN.

But also streams comprising IPN, which were not subjected to a prior imination step, can be directly fed into the first or first series of hydrogenation reactors.

Preferably, the hydrogenation of IPNI and/or IPN is conducted in the presence of ammonia. A suitable molar ratio of ammonia to IPNI and IPN in the hydrogenation step is about 5:1 to 500:1, preferably 10:1 to 400:1 and most preferably 20:1 to 300:1.

Preferably, ammonia is already added in the prior imination step. Additional ammonia can also be optionally added to bring the ammonia concentration into the aforementioned ranges.

The hydrogenation step is conducted in the presence of hydrogen.

The molar ratio between hydrogen and IPNI and IPN is preferably in the range of 3:1 to 10000:1, more preferably 4:1 to 5000:1 and most preferably 5:1 to 1000:1.

In a preferred embodiment, hydrogen is added after the imination step. It is however possible, that hydrogen is added prior to the imination step because the imination is usually carried out in the presence of catalysts which do not catalyse the hydrogenation of the imine or nitrile group.

The temperature during the hydrogenation is usually in the range of 40 to 200°C, preferably 50 to 150°C, more preferably 60 to 140°C and most preferably 60 to 130°C and a pressure of 50 to 300 bar, preferably 100 to 250 bar and more preferably 150 to 220 bar.

The hydrogenation is preferably carried out in the presence of hydrogenation catalysts, which usually comprise metals or semimetals from groups 1 to 17 of the Periodic Table as well as the rare earth metals.

Preferred catalyst elements are Ni, Co, Fe, Cu, Ru. Hydrogenation catalysts may also comprise Cr, Cu, Mo. Wo and/or Re.

Preferred hydrogenation catalysts comprise one or more of Ru and Co,

The hydrogenation catalysts can of the so-called Raney-type or the metal-oxide type.

Preferred Raney-type catalysts are Raney-Co-catalysts. The Raney-type catalysts may be supported or unsupported. Suitable Raney-Catalysts are further described in EP1207149, EP 2649042W02008107226, W02014086039 and WO2016120235, The hydrogenation catalysts can also be of the metal-oxide type.

Metal-oxide catalysts are preferably obtained by precipitation of soluble salts of the catalyst elements in the presence of catalyst supports to obtain the corresponding hydroxides, carbonates and oxides and which are usually transformed to the corresponding oxides during a calcination step. The precipitation step may also be conducted without the presence of support materials. Alternatively, the hydrogenation catalyst may be produced by impregnation of a catalyst support with soluble salts of the metals.

The metal-oxides catalysts are usually reduced in the presence of hydrogen prior to their use in the hydrogenation step. The reduced catalysts may be passivated by subjecting the reduced catalysts to an oxygen comprising gas in order to form a passivating and protective oxide layer which allows for safe handling and storage. The passivated catalysts may be reduced or activated prior to their use in the hydrogenation step. Activation and reduction of the metal oxide catalyst is preferably performed in the same reactor, in which the hydrogenation IPNI is performed. The reduction or passivation step may occur prior to the hydrogenation step, but it is also possible to reduce or activate the metal oxide catalysts in-situ during the hydrogenation of IPNI. The unreduced or inactivated catalyst is then transformed into its reduced form by the hydrogen present during the hydrogenation reaction.

Preferred supports are alumina, including but not limited to transitional alumina and non-tradi- tional alumina, titania, zirconia, silica, magnesia, calcium oxide and mixtures thereof.

In a further embodiment, the basicity of the feed comprising IPNI and/or IPN is increased prior or during the subsequent hydrogenation step.

In a preferred embodiment the basicity is increased after the first hydrogenation reactor or after the first series of hydrogenation reactors and before feeding the effluent from step b) or b-1) to the second hydrogenation reactor or the second series of hydrogenation reactors.

An increase of the basicity can be achieved by the addition of basic compounds or using hydrogenation catalysts which are supported on a basic support. Preferably the basic support comprises elements, such as oxides, of the alkaline metals, preferably Li, Na and K, the alkaline earth metals, preferably Mg and Ca or comprises basic minerals, preferably hydrotalcite, chrysotile or sepiolite.

Preferred basic catalysts are those which are disclosed in WO 2008077852.

In a most preferred embodiment, unsupported hydrogenation catalysts comprising 55 to 98 weight percent of Co, 0.2 to 15 weight percent of P, 0.2 to 15 weight percent of Mn and 0.2 to 15 weight percent of alkali, in particularly Na, are used. Details regarding the specification and production of such catalysts can be found in DE4325847.

Basic compounds can also be added in form of their solutions.

Suitable basic compounds are usually compounds of basic metals, in particularly the oxides, hydroxides or carbonates of alkaline metals, alkaline earth metals or the rare earth metals.

Other suitable basic compounds are ammonium hydroxide and amines.

Preferred basic compounds are oxides, hydroxides and carbonate, in particular U2O, Na2O, K₂O, Rb₂O, Cs₂O, LiOH, NaOH, KOH, RbOH, CsOH, Li₂CO3, Na2CO₃, K₂CO₃, Cs₂CO₃, Rb₂CO₃, MgO, CaO, SrO, BaO, Mg(OH)₂, Ca(OH)₂, Sr(OH)₂, Ba(OH)₂, MgCO₃, CaCO₃, SrCO₃ or

BaCO₃. In particularly preferred basic compounds are LiOH, NaOH and KOH.

The basic compounds are preferably added in form of their solutions in water or other suitable solvents, such as alkanols, like Ci-C4-alkanols, in particularly methanol or ethanol, or ethers, such as cyclic ether, in particularly THF or dioxane. Preferably the basic compounds are added in form of their aqueous solutions.

The concentration of basic compounds in water or other suitable solvents is usually around 0,01 bis 20 percent by weight, preferably 0, 1 bis 10 percent by weight and more preferably 0,2 bis 5 percent by weight.

The amount of added basic compound is usually determined in such a way as to yield a molar ratio of basic compound to IPNI and IPN is in the range of 100: 1 000 000 to 10 000: 1 000 000 and more preferably 200: 1 000 000 to 1000:1 000 000.

Further details regarding the addition of basic compounds prior to the hydrogenation step is disclosed in EP729937 or EP913387, whereas further details regarding the addition of basic compounds during the hydrogenation stage is disclosed in WO 2008077852.

The first hydrogenation reactor or the series of first hydrogenation reactions are preferably pressurized reaction vessels.

Most preferably each pressurized reaction vessel is a tubular reactor, and more preferably a tubular reactor where the hydrogenation catalyst is arranged in a fixed bed (fixed-bed reactor). Preferably, step a) is conducted in in 1 to 3, more preferably 1 to 2 and even more preferably in one single reactor.

If step a) is conducted in a series of two or more hydrogenation reactors, the feed preferably passes from the feed inlet of the first hydrogenation reactor to the effluent outlet of the last hydrogenation reactor in one pass with preferably no branching-off of any side stream. The catalyst load during the hydrogenation is preferably in the range of 0.01 to 10, preferably 0.03 to 5, more preferably 0.05 to 3 kg IPNI and IPNI per kg catalyst per hour.

According to the process of the present invention, at least a part of the effluent from the first hydrogenation reactor or the first series of hydrogenation reactors (step a) is fed to a second hydrogenation reactor or a second series of hydrogenation reactors (step b) and at least a part of the effluent from the first hydrogenation or the first series of hydrogenation reactors is recycled to the first hydrogenation reactor or the first series of hydrogenation reactors (step c).

The mass ratio of the part of the feed from step a) which is fed to step b) and the part of the feed from step a) which is recycled in step c) is preferably in the range of 0.1 :1 to 10:1 , more preferably 0.2:1 to 5:1 and most preferably 0.3:1 to 2:1.

In a particularly preferred embodiment the part of the effluent from step a) which is fed to step b) and the part which is recycled to step a) is controlled in such a manner that the cross-sectional loading of the first hydrogenation reactor is higher than the cross-sectional loading of the second hydrogen reactor or that at least the cross-sectional loading of at least one hydrogenation reactor in the first series of hydrogenation reactors is higher than the cross-sectional loading of at least one hydrogenation reactor of the second series of hydrogenation reactors. Preferably, the cross-sectional loading of all reactors of the first series of hydrogenation reactors is higher than the cross-sectional loading of all reactors of the second series of hydrogenation reactors.

In a preferred embodiment, the cross sectional loading of the first hydrogenation reactor or each reactor of the first series of hydrogenation reactors is 0.01kg/(m2s) or more, preferably 0.1 kg/(m2s) and more preferably 1 kg/(m2s) or more. In a more preferred embodiment, the cross sectional loading of the first hydrogenation reactor is 4 kg/(m2s) or more, preferably 5 kg/(m2s) and more preferably 7 kg/(m2s).

In a preferred embodiment, the cross sectional loading of the second hydrogenation reactor or each reactor of the second series of hydrogenation reactors is 4 kg/(m2s) or less, preferably 3 kg/(m2s) or less and preferably 2 kg/(m2s) or less.

In a more preferred embodiment, the cross sectional loading of the first hydrogenation reactor or each reactor of the first series of hydrogenation reactors is 4 kg/(m2s) or more, preferably 5 kg/(m2s) and more preferably 7 kg/(m2s) or more and the cross sectional loading of the second hydrogenation reactor or each reactor of the second series of hydrogenation reactors is 4 kg/(m2s) or less, preferably 3 kg/(m2s) or less and preferably 2 kg/(m2s) or less.

The cross-sectional loading of the reactor is calculated by means of formula (1) with the crosssection A of the reactor (unit [m2]) and the mass flow rate m (unit: [kg/s]) of liquid or dissolved reactants, ammonia and optionally solvents and/or further liquid reaction components (e.g., refluxes or liquid recycle streams). The gas phase of the reaction (e.g., hydrogen, inert gases) is not taken into account when calculating the cross-sectional loading.

In a preferred embodiment, the recycled effluent from step a) is cooled prior to feeding it back into the first hydrogenation reactor or the first series of hydrogenation reactors.

Preferably, recycle stream to step a) is cooled to a temperature in the range of 20 to 100°C, preferably 3 to 75°C and most preferably 40 to 60°C. Cooling the recycle stream has the advantage that the formation of hotspots in the first hydrogenation reactor or the first series of hydrogenation reactors in step a) may be avoided. Also, the CTR may be positively influenced. In a preferred embodiment, the process according to the present invention comprises an additional step a-1) after step a) in which the effluent from step a) is fed to a phase separator wherein the effluent from step a) is separated into a gas phase and a liquid phase.

In this embodiment, at least a part of the liquid phase from the separator is fed to step b), whereas the other part of the liquid phase is recycled to the first hydrogenation reactor, preferably to that part of the pipes or tubing through which the feed comprising IPNI and/or IPN is fed to step a). As specified above, the weight ratio of the part of the liquid phase fed to step b) and the part recycled to step a) is preferably in a range as to affect the preferred and most preferred cross-sectional loadings in the first hydrogenation reactor or the first series of hydrogenation reactors or the second hydrogenation reactor or the second series of hydrogenation reactors. The gas phase from the phase separator in step a-1) is preferably combined with the liquid phase from the phase separator which is fed to step b), i.e., to the second hydrogenation reactor or the second series of hydrogenation reactors. Alternatively, the gas phase is combined with the effluent form the second hydrogenation reactor or the second series of hydrogenation reactors. The separation step a-1) is preferably conducted at slightly lower pressures compared to the pressure of the first hydrogenation reactor or the first series of hydrogenation reactors. More preferably, the pressure difference between the first hydrogenation reactor or the first series of hydrogenation reactors and the second hydrogenation reactor or the second series of hydrogenation reactors is 1 to 10, preferably 2 to 8 and more preferably 3 to 7 bar. Removing a part of the gas phase in step a-1) has the advantage that less gas is liberated during the optional cooling step of the liquid phase obtained in step a-1) prior to its recycling to the first hydrogenation reactor or the first series of hydrogenation reactors.

Further, it is preferred to heat that part of the effluent from step a) or the liquid phase from step a-1) which is fed to step b). Heating is preferably carried out by a crossflow heat exchanger in which the heat of the effluent from step b) is used to heat the effluent from step a) or the liquid phase from step a-1). In this way, the heat of reaction generated in step b) is efficiently used to minimize the overall energy consumption of the process.

The second hydrogenation reactor or the second series of hydrogenation reactors are preferably pressurized reaction vessels.

If step b) is conducted in a series of two or more hydrogenation reactors, the feed preferably passes from the feed inlet of the first hydrogenation reactor of the second series of hydrogenation reactors to the effluent outlet of the last hydrogenation reactor in the second series of hydrogenation reactors in one pass, with preferably no branching-off of any side stream. The catalyst load during the hydrogenation in step b) is preferably in the range of 0.01 to 10, preferably 0.03 to 5, more preferably 0.05 to 3 kg IPNI and IPNI per kg catalyst per hour.

The reaction conditions, such as temperature, pressure and catalysts employed in step b), are substantially similar to the reaction conditions and catalysts used in step a), as specified above. Preferably, the temperature in the second hydrogenation reactor or the second series of hydrogenation reactors is 5 to 40°C, preferably 10 to 30°C above the temperature of the first hydrogenation reactor or the first series of hydrogenation reactors. The pressure of the second hydrogenation reactor or the second series of hydrogenation reactors is slightly lower than the pressure in the first hydrogenation reactor or the first series of hydrogenation reactors.

In a further preferred embodiment, a part of the effluent obtained from step b) is recycled to the first hydrogenation reactor or the first series of hydrogenation reactors or the second hydrogenation reactor or the second series of hydrogenation reactors. This embodiment has the advantage of providing further flexibility for finetuning the cross-sectional loading of each hydrogenation reactor or series of hydrogenation reactors to further optimize the yield and the selectivity for I PDA. Preferably, that part of the effluent from step b) which is recycled back to the step a) is cooled, preferably to a temperature in the range of 30 to 100°C, more preferably 40 to 85°C and most preferably 50 to 70°C. Cooling the recycle stream has the advantage that the formation of hotspots in the first hydrogenation reactor or the first series of hydrogenation reactors in step a) may be avoided. Also, the CTR may be positively influenced.

In a further preferred embodiment, the process according to the present invention comprises an additional step b-1) after step b) in which the effluent from the second hydrogenation reactor or the second series of hydrogenation reactors is fed to a phase separator wherein the effluent is separated into a gas phase and a liquid phase. Preferably, the effluent from step b) is cooled prior to feeding it into additional step b-1), as set out in the preceding paragraph.

The gas phase obtained in step b-1) is preferably recycled to step a) or step b) so that ammonia and/or hydrogen present in the gas phase can be directly used for the hydrogenation reaction in step a) or b). If the gas phase is directly recycled to steps a) and/or b), the subsequent refining section for the removal of hydrogen and/or ammonia from the effluent from step b) can be dimensioned smaller because the equipment will need to handle lower amounts of hydrogen and/or ammonia.

The effluent from step b) or step b-1), usually comprises cis-lPDA, trans-lPDA, IPNA, hydrogen, ammonia, components having a boiling point higher than cis-lPDA and optionally components having a boiling point lower than trans-lPDA.

When the effluent comprises hydrogen and ammonia, that part of the effluent from step b) or step b-1), which is not recycled, is usually worked-up by first separating hydrogen and ammonia.

The separation of hydrogen is preferably carried out by subjecting the effluent to a high-pressure separator which usually results in the separation of a gaseous phase, comprising hydrogen and some ammonia, and a liquid phase comprising ammonia, cis-lPDA, trans-lPDA, IPNA, components having a boiling point higher than IPNA and optionally, components having a boiling point lower than trans-lPDA.

The high-pressure separator is usually operated at pressure slightly lower than the pressure at which the hydrogenation reactor is operated, preferably of 2 to 350 bar, preferably 10 to 240 bar and more preferably 30 to 210 bar. The gaseous phase is preferably compressed to the reaction pressure and recycled to the hydrogenation reactor. The liquid phase from the high-pressure separator is usually subjected to one or more separation step, in which ammonia is separated from the rest of the components. Such separation steps may comprise one or more flash operations, stripping operations or distillation operations to obtain an ammonia fraction and the crude IPDA fraction. In a preferred embodiment, ammonia is separated in one or more distillation columns.

The distillation column is usually operated at pressures in the range of 5 to 50, preferably 10 to 40 and more preferably 15 to 30 bar. In a more preferred embodiment, a second ammonia removal step is conducted after the first ammonia removal step. Such a second step is preferably conducted in another distillation column usually operated at 1.5 to 20, preferably 2 to 15 and more preferably 3.5 to 10 bar.

The composition of the effluent from step b) after removal of ammonia and/or hydrogen is usually denoted as “crude IPDA”.

The purification of crude IPDA is known in the arts and is described in detail in PCT/EP2022/053216, W02004/020386, W02004/024668 and EP3275858.

Embodiments of the invention are illustrated in figures 1 to 3.

Figure 1 shows the simplest embodiment of the invention.

In Figure 1 a feed comprising IPNI and/or IPN is fed into a first hydrogenation reactor R1 in which at least a part of the IPNI and/or IPN is converted to IPDA. A part of the effuent from reactor R1 (step a) is recycled to the feed of reactor R1 .

The other part of the effluent from reactor R1 (step a) is fed into a second hydrogenation reactor R2. By splitting the effluent from step a) into a recycle stream and a feed stream to the second hydrogenation reactor R2, it is possible of operating both reactors with a different cross- sectional loading. By operating reactor R1 with a higher cross-sectional loading allows for the avoidance of hot spots leading to the formation of undesired side products. Operating reactor R2 with a lower cross-sectional loading allows to increase the conversion of IPNI and/or IPN. By recycling a part of the effluent from reactor R1 , the dimensions of reactor R2 may be reduced due to the lower feed rate. In a preferred embodiment the recycle stream is cooled, which allows for further reduction of hotspots in reactor R1. The part of the effluent from reactor R1 which is fed to reactor R2 may be heated prior to entering reactor R2 in order to increase the conversion rate in reactor R2. Preferably, some heat from the effluent of reactor R2 is used to heat the feed to reactor R2 in a cross-flow heat exchanger. It is also possible to recycle a part of the effluent from Step b) to the the feed of reactor R1 and/or to the reactor R2.

Figure 2 shows an embodiment of the invention comprising a first series of two hydrogenation reactor R1-1 and R1-2 (step a) and a second series of two hydrogenenation reactors R2-1 and R2-2 (step b).

A part of the effluent from the first series of hydrogenation reactors (R1-1 and R1-2) is recycled to the feed of reactor R1-1. Optionally, a part of the effluent from reactor R1-2 may also be recycled directly to reactor R1-2.

The other part of the effluent from the first series of hydrogenation reactors (R1-1 and R1-2) is fed into the first reactor R2-1 of the second series of hydrogenation reactors.

As described in Figure 1 , the recycle stream from reactor R1-1 may be cooled prior to feeding it into reactors R1-1 or R1-2. Also, the effluent from reactor R1-2 which is fed to reactor 2-1 may be heated, preferably with a cross-flow heat exchanger utilizing the heat of the effluent stream from reactor R2-2. Optionally, a part of the effluent from reactor R2-2 may also be recycled to reactors R1-1 or R1-2 or R-21 or R2-1 to allow for improved flexibility for finetuning the individual cross-sectional loadings to the individual reactors.

Figure 3 shows the embodiment shown and described in Figure 1 with an additional separator for carrying out step a-1) and another separator for carrying out step b-1). In step a-1), the effluent from reactor R1 is separated into a gas and a liquid phase. A part of the liquid phase is recycled to the feed of reactor R1 . The other part of the liquid phae is fed into reactor R2. The separated gas feed is reintroduced into the feed stream for reactor R2.

In step b-1), the effluent from reactor R2 is separated into a gas phase and a liquid phase. The gas phase is recycled to the feed of reactor R1 with the help of a compressor. The cooler at the outlet of reactor R2 is optional and allows to cool the effluent from reactor R2 prior to step b-1).

Surprisingly, it was found that the process according to the invention has the following advantages: the first hydrogenation reactor or the first series of hydrogenation reactors can be operated with a higher cross-sectional loading. This higher cross-sectional loading results in a more homogeneous temperature profile in the reactors. Hotspot formation may be significantly reduced. The avoidance of hotspots potentially results in fewer side products, increasing the selectivity and yield of IPDA. In addition, lower temperatures favor a high CTR, which is often desired in IPDA production. through the recycle stream, ammonia can be recycled to the reactor. In addition, the ammonia load to the downstream IPDA refining section, where hydrogen, ammonia and side products are separated can be reduced by the recycle stream. Further, the temperature of the recycle stream may be increased when two or two series of hydrogenation reactors are used. It was found that the hydrogen solubility was increased with increasing temperature. A higher hydrogen solubility in the liquid phase generally results in higher conversion rates and reduced side product formation. the second hydrogenation reactor or the second series of hydrogenation reactors can be operated with a lower cross-sectional loading. Thus, the residence time in the reactors can be increased, which generally leads to higher conversions. coupling a first hydrogenation reactor or a first series of hydrogenation reactors with a second or a second series of hydrogenation reactors according to the invention was also surprisingly found to reduce the overall energy duty required for the operation of the thermal equipment. the equipment for cooling and separating the liquid and the gas phase in the downstream process after the last hydrogenation step can be designed with smaller dimensions.

Claims

1 . A process for the manufacture of I PDA by hydrogenation of IPNI and/or IPN, comprising the steps of: a) feeding a stream comprising IPNI and/or IPN to a first hydrogenation reactor or a first series of hydrogenation reactors, b) feeding a stream comprising IPNI and/or IPN to a first hydrogenation reactor or a first series of hydrogenation reactors, c) recycling at least a part of the effluent from the first hydrogenation or the first series of hydrogenation reactors to the first hydrogenation reactor or the first series of hy drogenation reactors

2. A process according to claim 1 , wherein the first hydrogenation reactors or the series of first hydrogenation reactors and the second hydrogenation reactor or the series of second hydrogenation reactors are fixed bed reactors.

3. A process according to at least one of claims 1 to 2, wherein the hydrogenation reactors are tubular reactors.

4. A process according to at least one of claims 1 to 3, wherein the mass ratio of (i) the recycled part of the effluent of the first hydrogenation reactor or the first series of hydrogenation reactors to (ii) the part of the effluent fed to the second hydrogenation reactor or the second series of hydrogenation reactors is in the range of 0.1 :1 to 10:1.

5. A process according to at least one of claims 1 to 4, wherein cross sectional loading of the first hydrogenation reactor or each reactor of the first series of hydrogenation reactors is 4 kg/(m²s) or more.

6. A process according to at least one of claims 1 to 5, wherein cross sectional loading of the second hydrogenation reactor or each reactor of the second series of hydrogenation reactors is 4 kg/(m²s) or less.

7. A process according to at least one of claims 1 to 6, wherein the recycled effluent from step a) is cooled prior to feeding it back into the first hydrogenation reactor or the first series of hydrogenation reactors.

8. A process according to at least one of claims 1 to 7, comprising an additional step a-1) after step a) in which the effluent from the first hydrogenation reactor or the first series of hydrogenation reactors is fed to a phase separator wherein the effluent is separated into a gas phase and a liquid phase.

9. A process according to claim 8, wherein the liquid phase obtained in step a-1) is recycled to the first hydrogenation reactor or the first series of hydrogenation reactors.

10. A process according to at least one of claims 1 to 9, comprising an additional step b-1) after step b) in which the effluent from the second hydrogenation reactor or the second series of hydrogenation reactors is fed to a phase separator wherein the effluent is separated into a gas phase and a liquid phase.

11. A process according to at least one of claims 1 to 10 wherein at least a part the feed from step b) or at least part of the liquid phase obtained in step b-1) is recycled to the first hydrogenation reactor or the first series of hydrogenation reactors or the second hydrogenation reactor or the second series of hydrogenation reactors.

12. A process according to claim 10, wherein the gas phase obtained in step b-1) is recycled to the first hydrogenation reactor or the first series of hydrogenation reactors.

13. A process according to at least one of claims 1 to 12, comprising the additional step of cooling the effluent from the second hydrogenation reactor or the second series of hydrogenation reactors before step b-1).

14. A process according to at least one of claims 1 to 13, wherein the temperature of the first hydrogenation reactor or the first series of hydrogenation reactors is in the range of 80 to 140°C and the temperature of the second hydrogenation reactor or the second series of hydrogenation reactors is in the range of 80 to 150°C.

15. A process according to at least one of claims 1 to 15, wherein the pressure of the first hydrogenation reactor or the first series of hydrogenation reactors is in the range of 150 to 250 bar and the pressure of the second hydrogenation reactor or the second series of hydrogenation reactors is in the range of 150 to 250 bar.