WO1987007598A1

WO1987007598A1 - Process for hydrogenating an unsaturated organic compound

Info

Publication number: WO1987007598A1
Application number: PCT/GB1987/000340
Authority: WO
Inventors: George Edwin Harrison; Andrew George Hiles
Original assignee: Davy Mckee (London) Limited
Priority date: 1986-06-03
Filing date: 1987-05-18
Publication date: 1987-12-17
Also published as: EP0308411A1; ES2006770A6; CN87104061A; ZA873661B; DK51488D0; BR8707712A; IN169564B; AU7396187A; DK51488A; GB8613354D0; JPH01503454A

Abstract

A continuous process for hydrogenating an unsaturated organic compound to a corresponding hydrogenation product which process comprises: (a) providing a hydrogenation plant comprising first and second hydrogenation zones connected in series each containing a charge of a solid heterogeneous hydrogenation catalyst; (b) continuously supplying to an upper part of the first hydrogenation zone (i) a hydrogen-containing gas and (ii) a liquid phase containing the unsaturated organic compound dissolved in a compatible diluent therefor; (c) maintaining the first hydrogenation zone under temperature and pressure conditions conducive to hydrogenation; (d) allowing liquid phase to pass downwardly through the first hydrogenation zone; (e) continuously recovering an intermediate reaction product from a lower part of said first hydrogenation zone; (f) recovering a gaseous effluent from a lower part of the first hydrogenation zone; (g) supplying intermediate reaction product from step (e) in liquid form to an upper part of said second hydrogenation zone; (h) maintaining the second hydrogenation zone under temperature and pressure conditions conducive to hydrogenation; (i) allowing intermediate liquid reaction product to pass downwardly through said second hydrogenation zone; (j) supplying a hydrogen-containing gas to a lower part of the second hydrogenation zone; (k) recovering a gaseous effluent stream from an upper part of the second hydrogenation zone; (l) recovering a liquid hydrogenation product containing stream from a lower part of the second hydrogenation zone; and (m) purging material of at least one of the gaseous effluent streams of steps (f) and (k) from the hydrogenation plant; whereby the flows of gas and liquid are in co-current in said first hydrogenation zone and are in counter-current in said second hydrogenation zone.

Description

Process for hydrogenating an unsaturated organic compound.

This invention relates to a catalytic hydrogenation process. Heterogeneous catalytic hydrogenation processes of various kinds are widely practised on a commercial scale and are used for hydrogenation of a wide variety of unsaturated organic compounds. Typically such hydrogenation reactions are conducted at a pressure of from about 1 bar to about 300 bar and at a temperature in the range of from about 40°C to about 350°C Examples include hydrogenation of aldehydes to alcohols, of unsaturated hydrocarbons to saturated hydrocarbons, of acetylene-derived chemicals to saturated materials, of unsaturated fatty acids to saturated fatty acids, of ketones to secondary alcohols, of esters of unsaturated fatty acids to esters of partially or fully hydrogenated fatty acids, and of .certain sugars to polyhydroxyalcohols. Thus cyclqhexanol is produced commercially by catalytic hydrogenation of cyclohexanone, and iso-propanol by catalytic hydrogenation of acetone. An example of hydrogenation of an unsaturated hydrocarbon is the production of cyclohexane from benzene. Typical catalysts for such hydrogenation reactions include Group VIII metal catalysts, such as nickel, palladium and platinum.

Production of butane-l,4-diol by hydrogenation of but-2- yn-l,4-diol is an example of hydrogenation of an acetylene-derived chemical. A suitable catalyst for this reaction has been described as a granular nickel-copper- manganese on silica gel. The production of stearic acid by catalytic hydrogenation of the corresponding unsaturated acids, linoleic acid and linolenic acid, at a temperature of about 150°C and at a pressure of about 14.75 bar to about 32 bar and using a nickel, cobalt, platinum, palladium, chromium or zinc catalyst, is an example of the hydrogenation of unsaturated fatty acids to yield saturated fatty acids. So-called "hardening" of vegetable oils is an example of hydrogenation of esters of unsaturated fatty acids. As examples of hydrogenation of sugars to polyhydroxyalcohols there can be mentioned hydrogenation of aldoses to hexahydroxyalcohols, for example hydrogenation of D-glucose to sorbitol and of D- annose to mannitol.

An important route to C-, and higher alkanols involves hydroformylation of alpha-olefins, such as ethylene, propylene, and butene-1, to yield the corresponding aldehyde having one more carbon atom than the starting olefin. Thus ethylene yields propionaldehyde and propylene yields a mixture of n- and iso- butyraldehydes (with the nt-isomer usually predominating). These aldehydes yield the corresponding alkanols, ^"e.g. n- propanol and n_-butanόl, upon catalytic hydrogenation. The- important plasticiser alcohol, 2-ethylhexanol, is made by alkali-catalysed condensation of n-butyraldehyde to yield the unsaturated aldehyde, 2-ethyl-hex-2-enal, which is then hydrogenated to yield the desired 2-ethylhexanol. Although the preferred catalysts for such aldehyde hydrogenation reactions used to be Group VIII metal catalysts, such as nickel, palladium or platinum, the use of a solid catalyst comprising a reduced mixture of CuO and ZnO under vapour phase conditions has also been proposed (see EP-A-0008767 and US-A-2,549,416). Molybdenum sulphide supported on an activated carbon carrier has also been suggested in GB-A-765,972. The hydrogenation of an aldehyde feed containing ring-type sulphur compounds using a reduced mixture of oxides or hydroxides of copper and zinc is described in US-A- 4,052,467. Copper chromite has also been used as an aldehyde hydrogenation catalyst. Catalytic hydrogenation is in all the above cases a heterogeneous process. It may be operated as a liquid phase process or as a vapour phase process. A review of some of the factors involved in designing heterogeneous gas and vapour phase reaction systems appeared in "Chemical Engineering", July 1955, in an article entitled "Moving Bed - Processes ... New Applications", at pages 198 to 206 (see in particular pages 204 and 205 thereof).

There have been various prior proposals to operate hydrogenation proceses in several catalytic stages connected in series. For example, a vapour phase aldehyde hydrogenation process is described in US-A-4,451,677 which involves use of a plurality of adiabatically operated catalytic hydrogenation stages connected in series. In conventional multi-stage hydrogenation processes the hydrogen-containing gas and the material to be hydrogenated are fed through the plant in co-current or in counter-current fashion. In order to achieve good economy of hydrogen usage it is usual to recycle gas within the plant. Hence in designing the plant account must be taken of the circulating inert gases (e.g. N₂, Ar, CH, and the like) which are inevitably present in the circulating gas of a commercial plant.

The present invention seeks to provide an improved liquid phase hydrogenation process in which essentially 100% hydrogenation of the aldehyde or other unsaturated organic compound to the desired hydrogenation product can be achieved, without significant formation of byproducts. It further seeks to provide a multi-stage hydrogenation process in which the use of gas recycle compressors can be obviated. Additionally it seeks to provide a process for hydrogenation of a wide variety of unsaturated organic compounds which can be operated with excellent economy of hydrogen usage without the need for recycle of hydrogen-containing gases.

According to the present invention a continuous process for hydrogenating an unsaturated organic compound to a corresponding hydrogenation product comprises: (a) providing a hydrogenation plant comprising first and second hydrogenation zones connected in series each containing a charge of a solid heterogeneous hydrogenation catalyst;

(b) continuously supplying to an .upper part of the first hydrogenation zone (i) a hydrogen-containing gas and (ii) a liquid phase containing the unsaturated organic compound dissolved in a compatible diluent therefor;

(c) maintaining the first hydrogenation zone under temperature and pressure conditions conducive to hydrogenation;

(d) allowing liquid phase to pass downwardly through the first hydrogenation zone;

(e) ^' continuously recovering an intermediate reaction product from a lower part of the first hydrogenation zone;

(f) recovering a gaseous effluent stream from a lower part of the first hydrogenation zone;

(g) supplying intermediate reaction product from step (e) in liquid form to an upper part of the second hydrogenation zone;

(h) maintaining the second hydrogenation zone under temperature and pressure conditions conducive to hydrogenation;

(i) allowing intermediate liquid reaction product to pass downwardly through the second hydrogenation zone;

(j) supplying a hydrogen-containing gas to a lower part of the second hydrogenation zone;

(k) recovering a gaseous effluent stream from an upper part of the second hydrogenation zone; (1) recovering a liquid hydrogenation product containing stream from a lower part of the second hydrogenation zone; and

(m) purging material of at least one of the gaseous effluent streams of steps (f) and (k) from the hydrogenation plant; whereby the flows of gas and liquid are in co- current in the first hydrogenation zone and in counter- current in the second hydrogenation zone. The process of the invention is not specific to any particular hydrogenation reaction or to any particular catalyst composition. However, in general the hydrogenation conditions used in the first and second hydrogenation zones include use of a pressure of from about 1 bar to about 300 bar and of a temperature of from about 40°C to about 350°C.

The process of the invention can be applied, *for example to the hydrogenation of unsaturated hydrocarbons to saturated hydrocarbons. Typical of such a reaction is the production of cyclohexane from benzene. This hydrogenation can be carried out according to the invention using a nickel, palladium or platinum catalyst in each catalytic stage and a temperature of from about 100°C to about 350°C and a pressure of from ab'out 5 bar to about 30 bar. This reaction is exothermic. The use of high temperatures is normally recommended so as to maximise conversion of benzene to cyclohexane, but isomerisation of cyclohexane to methyl cyclopentane, which is extremely difficult to separate from cyclohexane, can occur in the aforementioned conventional procedures.

Production of secondary alcohols by reduction of ketones is another appropriate hydrogenation reaction to which the invention can be applied. Examples of such reactions include production of iso-propanol from acetone and of cyclohexanol from cyclohexanone. Another example of a hydrogenation reaction to ^* which the present invention can be applied is the production of butane-l,4-diol by hydrogenation of but-2- yn-l,4-diol. This can be carried out using a catalyst which is a granular nickel-copper-manganese on silica gel at a pressure of from about 200 bar to about 300 bar in each catalytic stage. A typical inlet temperature to the first hydrogenation zone is about 40°C, when the catalyst is freshly reduced. A further example of a hydrogenation reaction to which the process of the invention can be applied is the production of stearic acid by hydrogenation of linoleic acid, of lin^'olenic acid, or of a mixture thereof. This can be carried out using a nickel, cobalt, platinum, palladium, chromium or zinc catalyst at a pressure of from about 14.75 bar to about 32 bar and an inlet temperature to the first hydrogenation zone of about 150°C.

Other examples of hydrogenation processes to which the invention can be applied include "hardening" of vegetable oils and hydrogenation of sugars, (for example, hydrogenation of aldoses, such as D-glucose or D-mannose, to the. corresponding hexahydroxyalcohols, such as sorbitol and mannitol).

A particularly preferred type of hydrogenation reaction is the production of alcohols from aldehydes.

Such aldehydes generally contain from 2 to about 20 carbon atoms and may in the case of those aldehydes containing 3 or more carbon atoms include one or more unsaturated carbon-carbon bonds besides the unsaturated -CHO group. Thus as used herein the term "aldehyde" includes both saturated and unsaturated aldehydes, that is to say aldehydes wherein the only hydrogenatable group is the aldehyde group, -CHO, itself (such as alkanals) and aldehydes which contain further hydrogenatable groups such as olefinic groups, >C = C<, in addition to the - 1 -

aldehyde group, -CHO (such as alkenals). Typical aldehydes include acetaldehyde, propionaldehyde, n- and iso-butyraldehydes, ri-pentanal, 2-methylbutanal, 2- ethylhex-2-enal, 2-ethylhexanal, 4-_t-butoxybutyraldehyde, C-_j__Q-"OXO"-aldehydes (e.g. 2-propylhept-2-enal), undecanal, crotonaldehyde and furfural, as well as mixtures of two or more thereof. Examples of aldehyde hydrogenation reactions are the production of propanol from propionaldehyde, of n-butanol from n_-butyraldehyde, of 2- ethylhexanol from 2-ethylhex-2-enal, of undecanol from undecanal, and of 4-t-butoxybutanol from A-t- butoxybutyraldehyde. In such aldehyde hydrogenation reactions there can be used any of the conventionally used metal catalysts, such as Ni, Pd or Pt, or copper chromite, or a reduced mixture of CuO and ZnO of the type disclosed in EP-A-0008767 and US-A-2,549,416. According to EP-A-0008767 catalysts of this type under appropriately selected reaction conditions give rise to negligible formation of byproducts, such as ethers and hydrocarbons and also to small amounts only of "heavies" formation (such as esters) when aldehydes are hydrogenated.

Other aldehyde hydrogenation catalysts include cobalt compounds; nickel compounds which may contain small amounts of chromium or another promoter; mixtures of copper and nickel and/or chromium; and other Group VIII metal catalysts, such^" at Pt, Pd, Rh and mixtures thereof, on supports, such as carbon, silica, alumina and silica- alumina. The nickel compounds are generally deposited on support materials such as alumina or kieselguhr. The first and second hydrogenation zones may each include two or more beds of catalyst. Conveniently, however, each hydrogenation zone comprises a single catalyst bed. The individual beds of catalyst may be provided in separate vessels; in a preferred embodiment, however, the first and second hydrogenation zones comprise upper and lower beds respectively of catalyst housed within a single reaction vessel.

Thus in a preferred process the first and second hydrogenation zones comprise respective beds of catalyst mounted one above another within a reaction vessel and, which, in step (m), purging of material of the gaseous effluents of steps (f) and (k) is effected via a common purge gas system connected to the reaction vessel at a point or points between the beds of catalyst of the first and second hydrogenation zones. An alternative preferred process is one in which the first and second hydrogenation zones are provided in separate reaction vessels connected in series, in which the reaction vessel of the first hydrogenation zone is connected to the reaction vessel of the second hydrogenation zone by way of a conduit for liquid intermediate reaction product, and in which the reaction vessels of the first and second hydrogenation zones are each provided with a respective gas purge line for purging gaseous effluent therefrom. The hydrogen-containing gas supplied to the first hydrogenation zone preferably contains a major amount of hydrogen and at most a minor amount of one or more inert gases, such as nitrogen, methane, other low molecular weight hydrocarbons, such as ethane, propane, n- butane and iso-butane, carbon oxides, neon, argon or the like. Preferred hydrogen-containing gases are accordingly gases containing at least about 50 mole % up to about 95 mole % or more (e.g. about 99 mole %) of H₂ ^with the balance comprising one or more of N₂, CO, CO , Ar, Ne, CH_* and other low molecular weight hydrocarbons. Such hydrogen-containing gases can be obtained in conventional manner from synthesis gas and other usual sources of hydrogen-containing gases, followed by appropriate pre- treatment to remove impurities, such as sulphurous impurities (e.g. H₂S, CH₃SH, CH-_jSCH--, and CH₃SSCH₃) and halogen-containing impurities (e.g. CH--C1) which would exert a deleterious influence on catalytic activity, i.e. catalyst inhibition, poisoning or deactivation. Preparation of suitable hydrogen-containing gases will accordingly be effected according to usual production techniques and forms no part of the present invention.

The liquid phase supplied to the upper part of the first hydrogenation zone contains the unsaturated organic compound dissolved in a compatible diluent therefor. The purpose of the diluent is to act as a heat sink and to limit the temperature rise within the first hydrogenation zone to acceptable limits. The concentration of unsaturated organic compound in the liquid phase is accordingly preferably selected in dependence on the expected acceptable temperature rise across the first hydrogenation zone; such temperature rise should not be so great as to cause more than a minor- amount of vaporisation of the liquid ^"phase in the upper part of the first hydrogenation zone or to cause thermal damage to the catalyst. When the desired hydrogenation product and/or the diluent is relatively volatile, then it is possible to conduct the process so that a significant amount of vaporisation, or even complete vaporisation, occurs in the lower part of the first hydrogenation zone due to the adiabatic temperature rise caused by the heat released by the hydrogenation reaction. Such vaporisation is not deleterious to operation of the process, so long as the reaction mixture remains in the liquid phase in the upper part of the first hydrogenation zone. In this case the intermediate reaction product in the vapour phase exiting the first hydrogenolysis zone is desirably condensed for supply to the second hydrogenation zone in liquid form.

Generally speaking the liquid phase supplied to the first hydrogenation zone contains at least about 1 mole % of the unsaturated organic compound up to about 50 mole %, more preferably in the range of from about 5 mole % up to about 30 mole %, the balance being diluent or diluents. The diluent can be any convenient inert liquid or mixture of liquids that is compatible with the unsaturated organic compound, with any intermediate product or byproduct, and with the desired hydrogenation product. In many cases the hydrogenation product itself can be used as the compatible diluent or as a part of the compatible diluent. Hence, when hydrogenating an aldehyde, for example, the diluent can be the product alcohol obtained upon hydrogenation of the aldehyde. In this case the process of the invention includes the further step of recycling a part of the liquid hydrogenation product stream of step (1) for admixture with the unsaturated organic compound to form the liquid _r phase (ii) of step (b). Alternatively aldehyde condensation product, such as the dimers, trimers and higher condensation products of the type disclosed in GB- A-1338237, can be used as diluent. If the unsaturated organic compound used as starting material is a solid or if the hydrogenation product or an intermediate product is a solid, then an inert solvent will usually be used. Similarly, use of a solvent may be desirable in cases in which byproduct formation is a .problem. For example, hydrazobenzene is a potential byproduct of the hydrogenation of nitrobenzene to yield aniline; in such a case it is desirable to dissolve the unsaturated organic compound, such as nitrobenzene, in a solvent, such as ethanol, in order to limit formation of an undesirable byproduct, such as hydrazobenzene. In this case it is advantageous to include a minor amount of ammonia in the ethanol solvent as ammonia further limits the formation of byproducts such as azobenzene, azoxybenzene or hydroazobenzene.

The first hydrogenation zone may comprise an adiabatic reactor, a reactor with an internal cooling coil, or a shell and tube reactor. In the case of a shell and tube reactor the catalyst may be packed in the tubes with coolant passing through the shell or it may be the shell that is packed with catalyst with coolant flow through the tubes. The first hydrogenation zone is generally operated as a trickle bed reactor. In this case the hydrogen containing gas is generally admixed with the liquid phase upstream from the first hydrogenation zone and is partly dissolved therein. At the upper end of the first hydrogenation zone the concentration of unsaturated organic compound is at its highest in the liquid phase; hence the rate of hydrogenation is greatest at the upper end of the first hydrogenation zone. As the liquid phase passes downwar.dly through the first hydrogenation zone co- currently with the hydrogen it becomes depleted in respect of hydrogenatable material and to some extent in respect of dissolved hydrogen as the partial pressure of any inert gas or gases present rises and the partial pressure of hydrogen falls as the hydrogen is consumed by the chemical reactions taking place in the first hydrogenation zone. Hence at the lower end of the first hydrogenation zone the driving force for the .hydrogenation reaction is relatively low. The intermediate reaction product exiting the lower end of the first hydrogenation zone accordingly usually still contains a minor amount of chemically unsaturated hydrogenatable material. As already mentioned, the unsaturated organic compound used as starting material may include two or more hydrogenatable unsaturated groups which may undergo more or less selective hydrogenation in passage through the first hydrogenation zone. For^' example, when an olefinically unsaturated aldehyde (such as 2-ethylhex-2- enal) is hydrogenated, the olefinic bond tends to be hydrogenated first, before the aldehyde group, so that the saturated aldehyde (such as 2-ethylhexanal) is a recognisable intermediate product. However, some hydrogenation of the aldehyde group may occur prior to hydrogenation of the^' olefinic linkage, so that 2-ethylhex- 2-enol is an alternative intermediate product but is generally formed in lesser amounts. Each of these intermediates can then undergo hydrogenation to the desired alcohol product, e.g. 2-ethylhexanol.

When the unsaturated organic compound used as starting material contains only a single hydrogenatable material, then the unsaturated hydrogenatable organic material in the intermediate reaction product exiting the first hydrogenation zone will comprise the unsaturated organic compound itself. However, when the unsaturated organic compound used as starting material contains more than one hydrogenatable unsaturated linkage, then the unsaturated hydrogenatable organic material in the intermediate reaction product exiting the first hydrogenation zone will be selected from the starting material and any partially hydrogenated intermediates. For example, when hydrogenating 2-ethylhex-2-enal, the unsaturated organic material in the intermediate reaction product may be selected from 2-ethylhex-2-enal, 2- ethylhexanal, 2-ethylhex-2-enol, and a mixture of two or more thereof.

Generally speaking the hydrogenation conditions in the first hydrogenation zone are selected so as to effect hydrogenation of from^"about 75% to about 99% of the hydrogenatable unsaturated groups present in the unsaturated organic material supplied to the first hydrogenation zone. Typically the hydrogenation is completed to an extent of from about 85% to about 95% in the first hydrogenation zone. Such hydrogenation conditions include supply of hydrogen-containing gas to the upper part of the first hydrogenation zone in an amount sufficient to supply an amount of hydrogen that is greater than or equal to the stoichiometric quantity required to effect the desired degree of hydrogenation in the first hydrogenation zone. Usually it will be desirable to limit the supply of hydrogen-containing gas thereto so as to provide as nearly as possible such stoichiometric quantity of hydrogen and thereby to minimise hydrogen losses in the purge stream from the plant. Although the rate of supply of hydrogen-containing gas to the first hydrogenation zone will be to some extent dependent upon its composition, it will often be preferred to limit the rate of supply so as to provide not more than about 120% (e.g. up to about 110%), and even more preferably not more than about 105% (e.g. about 102%), of the stoichiometric quantity required to effect the desired degree of. hydrogenation in the first hydrogenation zone. The hydrogenation conditions will also be selected so that at least an upper part of the first hydrogenation zone is operated as a trickle bed reactor. Hence the rate of supply of the liquid feed will be limited by considerations such as the catalyst particle size and shape, the cross section of the reactor, and similar design factors.

The composition of the liquid feed will depend upon factors such as the exothermicity of the hydro¬ genation reaction, the maximum permissible temperature rise in the first hydrogenation zone, the design of the first hydrogenation zone, and the maximum permissible rate of supply to the first hydrogenation zone. When operating under adiabatic conditions the unsaturated organic compound (e.g. aldehyde):inert diluent molar ratio typically ranges from about 1:3 to about 1:10 and the rate of supply of liquid phase to the first hydrogenation zone ranges up to a rate corresponding to supply of unsaturated organic compound of about 8 moles per litre of catalyst per hour or more, e.g. up to about 10 or even 12 moles of aldehyde or other unsaturated organic compound per litre of catalyst per hour. If, however, provision is made for cooling the first hydrogenation zone as, for example, by use of internal cooling coils within the catalyst bed or by use of a shell and tube reactor, then a higher concentration of unsaturated organic compound can be used; hence in this case the unsaturated organic compound:inert diluent molar ratio typically ranges from about 1:1 up to about 1:10.

The inlet temperature to each of the hydrogenation zones will in each case be at least as high as the threshold temperature for the reaction and will be selected in dependence on the nature of the hydrogenation reaction. It will normally lie in the range of from about 40°C to about 350°C, whilst the operating pressure typically lies in the range of from about 1 bar to about 300 bar. For example when hydrogenating an aldehyde by the process of the invention the inlet temperature to the first hydrogenation zone is typically from about 90°C to about 220°C and the pressure is typically from about 5 to about 50 bar. Besides the unsaturated material and the hydrogenation product and diluent (if different from the hydrogenation product), the intermediate liquid reaction product leaving the first hydrogenation zone also contains dissolved inert gases and hydrogen. The gas phase leaving the first hydrogenation zone contains a higher level of inert gases than the hydrogen-containing gas supplied to the upper part of the first hydrogenation zone because hydrogen has been removed by the hydrogenation reaction in passage through the first hydrogenation zone. In the second hydrogenation zone the intermediate reaction product from the first hydrogenation ■ zone is fed in liquid form in counter-current to an upward flow of hydrogen-containing gas. The gas fed to the second hydrogenation zone may have the same composition as that supplied to the first hydrogenation zone. It is fed to the second hydrogenation zone generally in lesser amounts than the amount of hydrogen-containing gas supplied to the first hydrogenation zone. Generally speaking, it should be fed to the second hydrogenation zone in an amount sufficient to provide an at least stoichiometric amount of hydrogen corresponding to the amount of hydrogenatable material remaining in the intermediate liquid reaction product. Usually it will be preferred to supply hydrogen-containing gas to the second hydrogenation zone at a rate sufficient to supply not more than about 120% (e.g. up to about 110%), preferably not more than about 105% (e.g. about 102%), of the stoichiometric quantity of hydrogen required to complete the hydrogenation of the hydrogenatable organic material in the intermediate reaction product.

If desired, the gas fed to the second hydrogenation zone may be richer in hydrogen than that fed to the first hydrogenation zone. Hence the gas fed to the first hydrogenation zone may be, for example, a 3:1 molar ^H2^:N2 ^m;Lxture obtained by conventional methods from synthesis gas, whilst^" the hydrogen stream to the second hydrogenation zone is a substantially pure H₂ stream formed by subjecting the same H₂:N₂ mixture to purification e.g. by pressure swing absorption. In the second hydrogenation zone the highest H₂ partial pressure exists at the lower end thereof. Hence the driving force towards the desired hydrogenation product is maximised in the second hydrogenation zone and essentially all of the remaining unsaturated material in the intermediate reaction product exiting the first hydrogenation zone is hydrogenated in passage through the second hydrogenation zone.

An effluent stream comprising inert gases and hydrogen is taken from the plant between the first and second hydrogenation zones. This may be passed through a condenser in order to substantially recover any vaporised organic compounds therein. The resulting condensate is conveniently returned to the top of the second hydrogenation zone. The catalyst beds of the first and second hydrogenation zones will usually be supported on a suitable grid. When both beds are mounted in the same vessel, liquid intermediate reaction product from the first hydrogenation zone may simply be allowed to drop straight on top of the catalyst bed of the second hydrogenation zone. Usually, however, it will be desirable to collect and then to redistribute the liquid intermediate reaction product, evenly over the upper surface of the catalyst bed of the second hydrogenation zone with the aid of a suitable liquid distribution device. In some cases it may be desirable to collect and redistribute liquid within the first and/or second hydrogenation zones.

In a preferred process according to the invention for hydrogenation of an aldehyde the entry temperature to the first hydrogenation zone lies in the range of from about 90°C to about 220°C and the pressure lies in the range of from about 5 bar to about 50 bar. In operation of the process of the invention, under steady state conditions, the composition of the gas (whether dissolved in the liquid phase or present in the gaseous state) exhibits a significant variation between different parts of the plant. Thus, for example, the partial pressure of hydrogen is highest in each of the hydrogenation zones at the respective gas inlet end thereof and lowest at the exit end for gaseous effluent therefrom, whilst the combined partial pressures of any inert materials present is lowest at the respective gas inlet ends to the hydrogenation zones and highest at the exit ends for gaseous effluent therefrom. Under suitable operating conditions it is possible to operate the process of the invention so that the effluent gases contain a very small concentration of hydrogen (e.g. 5 mole % or less) and consist predominantly of inert gases (e.g. , Ar, CH^ etc). In this case the effluent gas stream or streams from the plant is or are relatively small and consequently hydrogen losses are minimal.

Because the inert gases are automatically concentrated in the gaseous effluent stream or streams, it is not necessary on economic grounds to recycle the gaseous effluents from the hydrogenation zones so as to obtain.efficient usage of hydrogen. Recycle of gas is necessary in conventional multi-stage co-current or counter-current hydrogenation processes in order to achieve efficiency of operation. Moreover, as it is not necessary to recycle a gas stream which contains appreciable concentrations of inert gases so as to achieve satisfactory economy of hydrogen consumption, the total operating pressure of the plant can be reduced; hence the construction costs can be reduced as the plant not only operates at a lower pressure but also no gas recycle compressor is needed. These factors have a significant effect on both capital and operating costs, both of which are lower for a plant constructed to operate the process of the invention than for conventional multi-stage co- current or counter-current hydrogenation plants.

In order that the invention may be clearly understood and readily carried into effect two preferred processes in accordance therewith will now be described, by way of example only, with reference to Figures 1 and 2 of the accompanying drawings, each of which is a simplified flow diagram of an aldehyde hydrogenation plant constructed in accordance with the invention.

It will be understood by those skilled in the art that Figures 1 and 2 are diagrammatic and that further items of equipment such as temperature and pressure sensors, pressure relief valves, control valves, level controllers and the like would additionally be required in a commerical plant. The provision of such ancillary items of equipment forms no part of the present invention and would be in accordance with conventional chemical engineering practice. Moreover it is not intended that the scope of the invention should be limited in any way by the precise methods of cooling and heating the various process streams, or by the arrangement of coolers, heaters, and heat exchangers, illustrated in Figures 1 and 2. Any other suitable arrangement of equipment fulfilling the requirements of the invention may be used in place of the illustrated equipment in accordance with normal chemical engineering techniques.

Referring to Figure 1 of the drawings, a stainless steel reactor 1 having an inside diameter of 6 inches (15.24 cm) and a height of 125 inches (317.5 cm) is provided with an upper stainless steel grid 2 which supports an upper bed 3 of a granular aldehyde hydrogenation catalyst. This catalyst is a prereduced nickel catalyst supported on 1/16 inch (1.6 mm) alumina spheres containing 61% of nickel (calculated as metal) in - the 50% reduced form and having a surface area of 140 m^/g as measured by the so-called BET method. The depth of bed 3 is 60 inches (152.4 cm) corresponding to a catalyst volume of 27.8 litres.

Reactor 1 is also fitted with a lower stainless steel grid 4 which supports a lower bed 5 of the same nickel catalyst. The depth of bed 5 is 30 inches (76.2 cm) corresponding to a catalyst volume of 13.9 litres. The distance between the top of lower bed 5 and upper grid 2 is 9 inches (22.9 cm). Thermocouples (not shown) are buried in catalyst beds 3 and 5 and reactor 1 is thermally insulated. Steam heating coils (not shown) are provided under the thermal insulation in order to assist in heating reactor 1 at start up.

The space 6 below lower grid 4 is 21 inches (53.3 cm) deep and is used to collect liquid emerging from the bottom of second bed 5. Such liquid is withdrawn by way of line 7 and is recycled by means of pump 8 and lines 9 and 10 through heat exchanger 11 and then through line 12 to a static liquid distributor 13 positioned above upper bed 3 at the top of reactor 1. Reference numeral 14 indicates a feed line for heat exchanger 11 for supply of a heating medium (e.g. steam) or cooling water as need arises. Heat exchanger 11 can be bypassed by means of bypass line 15, flow through which is controlled by means of a valve 16 coupled to a temperature controller 17 which monitors the temperature in line 12. Aldehyde to be hydrogenated is supplied in line 18 and admixed with the liquid exiting heat exchanger 11. Alcohol hydrogenation product is withdrawn by way of line 19 under the control of valve 20 which is itself controlled by means of a level controller 21 arranged to monitor the liquid level in bottom space 6 of reactor 1.

Hydrogen-containing gas is supplied to reactor 1 in line 22. A major part of the gas flows in line 23 to the top of reactor 1 under the control of a flow controller 24 whilst the remainder is fed by way of line 25 under the control of a further flow controller 26 to an upper part of the bottom space 6 at a point above the liquid level in bottom space 6.

A gas purge stream is taken from the space 27 between the two catalyst beds 3 and 5 in line 28. This is passed through a condenser 29 supplied with cooling water in line 30. Condensate is collected in drum 31 and is returned to reactor 1 in line 32. The resulting purge gas stream is taken in line 33 and passed through a further condenser 34 which is supplied with refrigerant in line 35. Pressure control valve 36 is used to control the pressure within the apparatus and the rate of withdrawal of purge gas in line 37.

Reference numeral 38 indicates a static liquid distributor for distributing evenly across the top of lower bed 5 liquid that exits upper bed 3. Line 39 and valve 40 are used for initial charging of th.e reactor 1 with liquid.

Reference numeral 41 indicates an optional internal cooling coil which is supplied with cooling water in line 42.

The plant of Figure 2 is generally similar to that of Figure 1 and like reference numerals have been used therein to indicate like features. Instead of a single reactor vessel 1 the plant of Figure 2 has two separate reactors 43, 44 each containing a respective catalyst bed 3, 5. Liquid intermediate reaction product emerging from the bottom of first catalyst bed 3 collects in the bottom of reactor 43 and passes by way of line 45 to the top of reactor 44. Purge gas is taken fro^"m reactor 43 in line 46 and from reactor 44 in line 47 which joins line 46 to form line 48 which leads in turn to condenser 29. Condensate is returned via line 32 from drum 31 to the top of reactor 44.

The apparatus of Figure 2 permits operation of the two reactors 43 and 44 at different pressures; in this case a valve (not shown) can be provided in one or both of lines 46 and 47 and a pump (not shown) can be provided, if necessary, in line 32. The invention is further illustrated in the following Example. Example

After loading reactor 1 with catalyst, the apparatus of Figure 1 is purged with nitrogen and then pressurised with nitrogen in order to test for leaks. Reactor 1 is then charged with undecanol by way of line 39. Pump 8 is then started and circulates undecanol through lines 10 and 12 at a rate of 119.6 kg/h. A hydrogen containing gas with the composition 95% v/v H₂, 4% v/v CH₄ and 1% v/v N₂ is supplied at a low rate through line 22 and the system is purged by way of line 37. Valve 36 is then set to control the pressure within the apparatus at 250 psig (18.24 bar). Steam is supplied to heater 11 in line 14 to raise the temperature of reactor 1 and its contents to just above 115°C; steam is also supplied to the steam heating coils (not shown) wound around the external wall of reactor 1 under the layer o^'f thermal insulation in order to hasten the warm up phase at start up of the apparatus.

A flow of 35.505 kg/h of undecanal at 20°C is established in line 18 and gas flows of 5.165 Nm^/h are established through flow controller 24 and of 0.246 Nm^/h through flow controller 26, using a number of step changes. Heat exchanger 11 is then fed with cooling water in line 14 and temperature controller 17 is set to 115°C. The temperature of liquid exiting heat exchanger 11 is 143.2°C.

When all flows are established and the apparatus has stabilised thermally, liquid accumulates in bottom space 6 at a temperature of about 150°C. Product is withdrawn in line 19 at a rate of 35.925 kg/h under the control of level controller 21.

Analysis of the cooled product in line 19 by gas-liquid chromatography shows that it contains more than 99.8 molar %^" undecanol (i.e. predominantly undecan-1-ol with a trace of 2-methyldecan-l-ol) and less than 0.2 molar % heavy byproducts derived by self-condensation of the aldehydes in the feed stream. This analysis also shows that less than 50 ppm of carbonyl compounds

(undecanal) remain in the hydrogenated product. This analysis neglects the small amount of disolved gases in the product stream.

Approximately 0.73 Nm³/h of gas containing 63.3% v/v of hydrogen and 36.7% v/v of methane plus nitrogen is discharged from the apparatus in line 37. A small amount of organic vapours in the gas in line 28 is condensed in condenser 29 and runs back down into reactor 1 by way of drum 31 and line 32. Temperatures, pressures, and stream compositions remain essentially constant during a 10 day run indicating stable catalyst activity. The proportion of heavy aldehyde condensation byproducts decreases slightly during the course of the run; during the filial 3 days of the run the level of byproducts is less than 0.1% molar. The plant of Figure 1 and the operating techniques described above are generally applicable to hydrogenation of organic materials. It will accordingly be readily apparent to the skilled reader that the teachings of the invention can be practised with a wide variety of hydrogenation reactions other than the aldehyde hydrogenation reaction specifically described in relation to Figure 1 of the accompanying drawings.

Claims

1. A continuous process for hydrogenating an unsaturated organic compound to a corresponding hydrogenation product which process comprises:

(a) providing a hydrogenation plant comprising first and second hydrogenation zones connected in series each containing a charge of a solid heterogeneous hydrogenation catalyst; (b) continuously supplying to an upper part of the first hydrogenation zone (i) a hydrogen-containing gas and (ii) a liquid phase containing the unsaturated organic compound dissolved in a compatible diluent therefor;

(e) continuously recovering an intermediate reaction product from a lower part of said first hydrogenation zone;

(f) recovering a gaseous effluent from a lower part of the first hydrogenation zone;

(g) supplying intermediate reaction product from step (e) in liquid form to an upper part of said second hydrogenation zone;

(h) maintaining the second hydrogenation zone under temperature and pressure conditions conducive to hydrogenation; (i) allowing intermediate liquid reaction product to pass downwardly through said second hydrogenation zone;

(j) supplying a hydrogen-containing gas to a lower part of the second hydrogenation zone; (k) recovering a gaseous effluent stream from an upper part of the second hydrogenation zone;

(1) recovering a liquid hydrogenation product containing stream from a lower part of the second hydrogenation zone; and (m) purging material of at least one of the gaseous effluent streams of steps (f) and (k) from the hydrogenation plant; whereby the flows of gas and liquid are in co- current in said first hydrogenation zone and are in counter-current in said second hydrogenation zone.

2. A process according to claim 1, in which the first and second hydrogenation zones comprise respective beds of catalyst mounted one above another within a reaction vessel and in which, in step (m), purging of material of the gaseous effluent of steps (f) and (k) effected via. a common purge gas system c nnected tp the^* reaction vessel at a point or points intermediate the beds of catalyst of the first and second hydrogenation zones.

3. A process according to claim 1, in which the first and second hydrogenation zones are provided in separate reaction vessels connected in series, in which

9 the reaction vessel of the first hydrogenation zone is connected to the reaction vessel of the second hydrogenation zone byway of a conduit for liquid intermediate reaction product, and in which the reaction vessels of the first and second hydrogenation zones are each provided with a respective gas purge line for purging gaseous effluent therefrom.

4. A process according to any one of claims 1 to 3, in which the hydrogen-containing gas of step (b) comprises a major molar amount of hydrogen and a minor molar amount of one or more inert gases.

5. A process according to any one of claims 1 to 4, in which the hydrogen-containing gas of step (j) is richer in hydrogen than that of step (b).

6. A process according to any of claims 1 to 5, in which the compatible diluent comprises said hydrogenation product.

7. A process according to claim 6, which includes the further step of recycling a part of the liquid hydrogenation product stream of step (1) for admixture with the unsaturated organic compound.to form the liquid phase (ii) of step (b).

8. A process according to any one of claims 1 to 7, in which the unsaturated,organic compound is an aldehyde, and in which the hydrogenation product is an alcohol.

9. A process according to claim 8, in which the aldehyde is propionaldehyde, and in which the alcohol is ri-propanol.

10. A process according to claim 8, in which the aldehyde is ri-butyraldehyde, and in which the alcohol is ri-butanol.

11. A process according to claim 8, in which the aldehyde is 2-ethylhex-2-enal, and in which the alcohol is 2-ethylhexanol.

12. A process according to claim 8, in which the aldehyde is undecanal, and in which the alcohol is undecanol.

13. A process according to any one of claims 8 to 12, in which the entry temperature to the first hydrogenation zone lies in the range of from about 90°C to about 220°C and in which the pressure lies in the range of from about 5 to about 50 bar.