US3202723A

US3202723A - Process for the catalytic hydrogenation of aromatic hydrocarbons

Info

Publication number: US3202723A
Application number: US223393A
Authority: US
Inventors: Thonon Clement
Original assignee: IFP Energies Nouvelles IFPEN
Current assignee: IFP Energies Nouvelles IFPEN
Priority date: 1961-09-13
Filing date: 1962-09-13
Publication date: 1965-08-24
Anticipated expiration: 1982-08-24
Also published as: GB955926A; FR1419557A; OA01615A

Description

Aug 24,

HYDROGEN 6 QwuMP c. THONON 3,202,723 PROCESS FOR THE CATALYTIC HYDROGENATION OF AROMA'IIC HYDROCARBONS Filed Sept. 15, 1962 J 24 I L LHYDROGEN RECYCLE H 5 SECOND REACTION ZONE 5 5 cowoamszn 5 15 16 SEPARATING UNIT 14 9PUMP 4 CYCLOALIPHATIC HYDROCARBON FIRST REACTION ZONE HEAT EXCHANGER PUMP AROMATIC HYDROCARBON INVENTOR CLEMENT THONO/V ATTORNEY$ 3,202,723 PROCESS FOR THE CATALYTIC HYDROGENA- TRON F AROMATIC RQCARBONS Clement Thonon, Seine-et-Oise, France, assignor to Invstitut Francais (in Petrole des 'Carhurants et Lubrifiants,

Rueil-Malmaison,Seine-et-Oise, France Filed Sept. 13, 1962, Ser. No. 223,393 Qlaims priority, applicatiorgrance, Sept. 13,1961,

9 Claims. hi. 260-667) version is conducted with the starting material in the vapor phase. These known processes, however, suffer from various disadvantages, particularly if it is desired to produce a cycloaliphatic hydrocarbon of a high purity as presently required for industrial uses.

The vapor phase processes yield an undesirably low output per unit volume of reaction zone. This is due not only to the low density of the treated product but also to the difficulties encountered in attempting to cool said reaction zone efiiciently. It is necessary either to use a bulky apparatus comprising critical and costly internal cooling circuits so as to completely avoid local overheating of the catalyst which if uncontrolled would result in a reduction of catalyst activity; or, on the other hand, to dilute the aromatic hydrocarbons in a large volume of gases or vapors, for instance, in the corresponding hydrogenated products, which requires the recycling of a large portion of the latter and accordingly leads to a poor hourly rate of production of the plant.

Considering now the liquid phase processes, particularly those in which the aromatic hydrocarbons, together with hydrogen, are introduced into a suspension of the catalyst in a large excess of the corresponding cycloaliphatic hydrocarbon, it is to be noted that although these systems provide for an eiiicient removal of the heat produced by the reaction (particularly when the hydrogenated hydrocarbon is removed by distillation) these systems are far from entirely satisfactory, primarily because are products obtained thereby are not of a sufficiently high purity. As a matter of fact, it is not technologically practical during such a distillation, to completely separate treated hydrocarbon from the hydrogenated product, as these two products have very close boiling temperatures. Furthermore, if an etiort is made to convert substantiallyall the aromatic hydrocarbon in the reaction zone, the rate of reaction is substantially decreased under these conditions.

Accordingly, the liquid phase processes are not easily applicable when it is desired to get a large output of cycloaliphatic hydrocarbons of a high purity, since such a result is only achieved by use of either of a very bulky apparatus or of a series of reaction vessels of substantially equal volumes in the first of which the major part (for instance 95%) of the aromatic hydrocarbon is converted, whereas in the other reaction vessels the conversion is limited to decreasing portions of the aromatic hydrocarbons (for instance 3.5% in a second reaction vessel, 1% in a third one, 0.3% in a fourth one, etc.), the cycloaliphatic hydrocarbon produced being eventua United States Patent reaction.

ally separated from the reaction product obtained at the outlet of each of the successive reaction vessels, by dis tillation. Such a process requires the use of a very large and expensive equipment as well as the addition of external for carrying out the separation of the cycloaliphatic hydrocarbon by distillation in the last stage or stages of As a matter of fact, in view of the low conversion rate in these last stages, the heat generated by the reaction is neglible and, therefore, insuflicient for the distillation of the reaction product. It is also impossible to use the heat generated by the reaction in the preceding stages for distillation of the last stage products since the heat is produced at too low'a temperature level.

It is, therefore, an object of this invention to provide for the converison of aromatic hydrocarbons, such as, for instance, benezene, toluene and/or xylene, to the corresponding cycloaliphatic hydrocarbons, under such conditions as to obtain the latter with a high degree of purity.

It is another object of this invention to carry out said conversion with a suifieient reaction rate per unit volume of the reaction zone.

It is yet another object of this invention to avoid the use of extensive and expensive cooling means in the reaction zone and/or local overheating of the catalyst.

It is still another object of this invention ,to carry out the conversion of aromatic hydrocarbons to the corresponding cycloaliphatic hydrocarbons without dilution of said aromatic hydrocarbons in the vapor phase with a large volume of gases or vapors which have to be recycled.

It is a further object of this invention to carry out catalytic hydrogenation of aromatic hydrocarbons with high conversion rates at a satisfactory reaction velocity.

It is still a further object of this invention to provide for the conversion of aromatic hydrocarbons to the corresponding cycloaliphatic hydrocarbons without use of substantial amounts of external heat.

These and other objects as may be apparent from a study of the following specification and claims, are achieved by the process of this invention as follows:

(1) The aromatic hydrocarbon is introduced together with hydrogen in excess or in amount at least a stoichiometric ratio, in a first reaction zone containing the corresponding cycloaliphatic hydrocarbon in the liquid phase having suspended therein a catalyst for liquid phase hydrogenation, the molar ratio of said aromatic hydrocarbon to the corresponding cycloaliphatic hydrocarbon being maintained at a substantially constant value comprised between 0.003 and 0.1; and,

(2) The gaseous flow issuing from the bath is then caused to pass through a second reaction zone containing a solid catalyst for hydrogenation in the gaseous phase, the outlet gaseous flow issuing from said second reaction zone consisting of excess hydrogen and the cycloaliphatic hydrocarbon of high purity which may then be condensed by cooling.

. The attached drawing depicts in fiowsheet form, an embodiment of this invention.

Under the conditions of this invention, a cycloaliphatic hydrocarbon of high purity is obtained with a high output rate. Furthermore, the heat necessary for vaporizing the hydrocarbons in view of their treatment in the second reaction zone, is only that heat generated by the reaction conducted in the first reaction zone, so that any addition of external heat may be avoided.

The molar ratio between the aromatic hydrocarbon and the cycloaliphatic hydrocarbon may be easily determined by conventional means, for instance, by ultraviolet spectrography and gas chromatography. In the first reaction zone, the molar ratio of aromatic to cycloaliphatic hydrocarbons is 0.003-0.l:l, preferably (LOGS-0.05:1

Patented Aug. 24, 1965 respectively. The intermediate product stream which is fed to the gas phase reaction zone generally contains the aromatic to cycloaliphatic hydrocarbons in a molar ratio of about 0.003-0.1:1, preferably 0005-005: 1, respectively.

The catalysts used in each of the two stages of reaction are conventional hydrogenation catalysts. The catalyst in the first reaction zone is a solid conventional catalyst for liquid phase hydrogenation, and it is suspended in the liquid hydrocarbon. In the second reaction zone, a conventional catalyst for gaseous phase hydrogenation may be used in the form of a stationary, a moving or a fluid bed, the stationary bed being, however, preferred. These catalysts may be used either as such, or deposited on any carrier.

By way of example, there may be used as a catalyst, a metal of the 8th group of Mendeleefis periodic table; particularly nickel, cobalt, platinum, palladium, rhodium, or ruthenium, and as a support there may be used alumina, silica, pumice stone, asbestos, clays and the like.

The reaction temperature and pressure are selected in the range of the temperatures and pressures commonly used when operating with the previously mentioned hydrogenation catalysts, and at such a level within said range as to maintain a liquid phase during the first step of the process and a gaseous phase during the second step.

In the first reaction zone the operating temperature will preferably be about 10 C. to 100 C., preferably about 30 C. to 70 C. below the actual boiling temperature of the cycloaliphatic hydrocarbon under the prevailing total pressure. By way of example, satisfactory reaction velocities for an industrial plant may be obtained at temperatures between 80 and 250 C. and under pressures in the range of from 1 to 100 atmospheres, these limits being not, however, strictly obligatory. The temperature and pressure used for the first step may be difierent from those selected for conducting the second step.

The two steps may be, if desired, conducted under substantially the same temperature and pressure conditions, provided that the partial pressure of the hydrocarbons in the outflow of the first stage he reduced, for instance by dilution of the same with hydrogen or recycled gas so as to avoid any undesirable condensation of the hydrocarbons during the second stage of the reaction. Alternatively, condensation may be prevented by operating the second reaction zone at a lower pressure or higher temperature than the first zone.

The amount of hydrogen used is at least equal to the stoichiometric proportion corresponding to the desired degree of saturation, for instance 3 mols of hydrogen per mol of benzene, toluene or xylene and 2 or 5 mols of hydrogen per mol of naphthalene, etc. An excess of 100% is preferred. Hydrogen may be used either in pure form or in admixture with other diluent gases such as, for instance, methane, or nitrogen. Furthermore, it is of advantage to recycle the excess hydrogen separated from the outflow of the second reaction zone, said hydrogen stream containing significant proportions of such diluent gases.

This invention will be further explained in more detail with reference to the accompanying drawing showing a fiowsheet of a preferred embodiment of the present invention.

The schematically illustrated apparatus comprises essentially, a first reactor 1 wherein is conducted the first hydrogenation step, a heat exchanger 2, a second reactor 3, a condenser 4, a separating unit 5 for separation of incondensable gases from the liquid product, and pump means 6, 7, 8 and 9.

The reaction vessel 1 containing the cycloaliphatic hydrocarbon having the catalyst suspended therein, is fed with the corresponding aromatic hydrocarbon through the pipe 10, the circulating pump 7, and the pipe 11. Hydrogen is conveyed through pipe 12, the circulating benzene.

At the outlet of said condenser the outflow is partially condensed and caused to pass through pipe 17 to the separating unit 5 from where the pure liquid cycloaliphatic hydrocarbon is removed through pipe 18 whereas the incondensable gases are recycled to the input of the first reaction vessel through pipe 19.

A part of the reaction mixture contained in the reaction vessel 1 is continuously withdrawn through pipe 20, passes through the pump 8, the pipe 21 and the heat exchanger 2 and is recycled to the reaction vessel 1 so as to maintain the temperature therein substantially constant. It is to be appreciated, however, that other circulation systems for removal of heat may be used such as thermo-siphon cooling, heat exchange means internal to the reaction vessel Or the like.

According to another embodiment of apparatus for carrying out the process of the invention, the reaction vessel 3 may be juxtaposed to the upper part of the reaction vessel 1.

Various modifications may be made to the immediately describe-d system. Thus, for instance, a portion of the gases recycled through pipe 19 may be conveyed through pipe 23 to the reaction vessel 3, said gases constituting a diluent which prevents any undesirable condensation of hydrocarbon vapors in the reaction vessel 3. Another improvement consists of withdrawing from the pipe 24 a part of the recycled gases so as to avoid any excessive accumulation of inert gases such as CH CO, N and the like.

Alternatively, another embodiment which is not shown may consist of liquefying a portion of the recycled gases so as to separate therefrom the inert gases which are more easily liquefiable than hydrogen. Obviously, other design and engineering changes may be made by those skilled in the art without departing from the essence'of the invention. To demonstrate the technological progress and unobviousness of the present invention, the following comparative examples illustrate the advantages of the present process (Examples 1, 6 and 7) as compared to the prior art processes (Examples 25).

Example 1 A conversion is conducted in an apparatus corresponding to the fiowsheet of FIGURE 1, at a temperature of 200 C. and under a. pressure of 40 kg./cm. in both reaction vessels 1 and 3.

Forty-five kg. of cyclohexane are introduced together with 5 kg. of divided Raney nickel into the reaction vessel 1 which is thereafter fed with benzene at a rate of kg./hour and with hydrogen in 30 %excess of the stoichiometric ratio. Under these conditions, the reaction volume is kept constant, with a molar ratio benzene/cyclo hexene in the liquid and gaseous phases of about 0.015. The gaseous outflow from reaction vessel 1 is then passed through the reaction vessel 3 containing 2.5 kg. of a catalyst consisting of activated carbon having a 20% nickel content deposited thereon, said catalyst being used in the form of a stationary bed.

The cyclohexane product is obtained in a substantially quantitative yield by condensation of the effluent from the reaction vessel 3 and the product contains only 0.01% Furthermore, no additional heat is required during the reaction.

Example 2 'the liquid phase in a single reaction vessel 1. It is thus 1500% increase in the undesired impurity.

Example 3 Example 2 is repeated with a reaction'mass consisting of 20 kg. of R-aney nickel and 180 kg. of cyclohexane. Under these conditions, the cyclohexane obtained as reaction product still contains 0.12% benzene, which is equivalent to a 230% increase in the undesired impurity.

' Example 4 Example 1 is repeated except that 'thereaction vessel 3 is replaced by another reaction vessel for liquid phase operation which is identical to the reaction vessel 1, each of the two reaction vessels initially containing- 5 kg. of Raney nickel and 45 kg. of cyclohexane.

Under these conditions (two reaction vessels for liquid phase operation) the cyclohexane obtained as reaction product still contains about 0.1% benzene. Consequent- 1y, this example not only yields about a 100% increase in impurities as compared to Example 1, but it also requires the use of a large amount of external heat for thesecond reaction vessel so as to secure the removal of the reaction product inthe vapor phase. Thus, this example yields a less pure product as well as a more expensive operation.

Example 5 According to this example, there is used only one reaction vessel for operation in the vapor phase. In order to secure a conversion of 100 kg./hour of benzene, it is necessary to make use of a reaction vessel having a volume corresponding to about 20 times the volume of the reaction vessel 3 for operation in the vapor phase according to Example 1, and of 50 kg. of a catalyst consisting of nickel deposited on active carbon and used in the form of a stationary bed. It is also necessary to recycle and cool the eflluent flow from the reaction vessel at a rate of 1500 kg./hour so as to remove the heat generated by the reaction. 7

Under such conditions it is impossible to maintain a constant temperature in the reaction vessel since the temperature at the feed thereof is 160 C., whereas at the outlet the temperature is 220 C.

In addition to these drawbacks of great equipment expense and lack of control, the cyclohexane obtained as reaction product still contains a 0.09% benzene content, which amounts to an 80% increase in the content of the undesired impurity, which is of course again highly dele terious.

The preceding five examples repeated except with different hydrogenation catalysts and/or diflerent temperatures or pressures as taught previously, yield substantially the same comparative results. It is to be emphasized that the essence of this invention resides in a liquid phase catalyst hydrogenation reaction followed immediately by a gas phase hydrogenation catalytic reaction to result in an economical and simple process yielding highly pure cycloparaffins substantially devoid of the aromatic starting materials.

It will be understood that while there have been given herein certain specific examples and suggestions for the practice of this invention, it is not intended thereby to have the invention limited to or circumscribed by the specific details ofcatalyst starting materials, proportions or operating conditions herein specified, in view of the fact that the invention may be modified according to individual preference or conditions without departing from the spirit and scope of this disclosure and thereby being within the range of equivalence of the following appended claims.

Such modifications will be exemplified by the following non-limitative examples. 7 V

Example 6 Example 1 is repeated, except that there is initially introduced into the reactionvessel ,1, 67.5 kg. of'cyclohexane and 75 kg. of a divided catalyst consisting of 35% by weight kieselguhr and 65% by weight nickel deposited? thereon, and except that the reaction vessel 3 contains. 3.5 kg. of a catalyst consisting of 15% by weight. nickel and 85% by Weight alumina, said catalyst having been previously activated at a temperature of about 340". C. for 6 hours in hydrogen.

The results of Example 1 remain substantially unchanged. I

Example 7 Into an apparatus corresponding to theflowsheet of FIGURE .1, there is introduced 90 kg. of methylcyclohexane and 10 kg. of divided Raney nickel, and then toluene at a rate of 100 kg./hour and hydrogen in excess over the stoichiometric ratio.

The gaseous outflow is then passed through the reaction vessel. 3 containing-4 kg. of .a catalyst consisting of 20% by weight nickel and 80% by weight activated carbon, said catalyst being used in the form of a stationary bed. Methylcyclohexane containing 0.03% benzene is obtained in a substantially What is claimed is:

1. A process for the catalytic, non-destructive hydrogenation of at least one aromatic hydrocarbon, which process comprises the steps of:

(1) introducing said aromatic hydrocarbon in the liquid state into the corresponding liquid cycloaliphatic hydrocarbon contained in a first reaction zone, passing through said cycloaliphatic hydrocarbon, an amount of gaseous hydrogen at least equal to the stoichiometric ratio with respect to the aromatic hydrocarbon, said liquid cycloaliphatic hydrocarbon having a Group VIII metal hydrogenation catalyst suspended thereinto, the molar ratio between said aromatic hydrocarbon and said cycloaliphatic hydrocarbon in said first reaction zone being maintained within the range of from 0.003:l to 0.111, the pressure being comprised between about 1 and 100 atmospheres and the temperature comprised between and 250 C., provided the temperature is at least 10 C. lower than the initial boiling temperature of said cycloaliphatic hydrocarbon under the prevailing pressure, and

(2) passing as such the whole, unreacted hydrogen and vaporized hydrocarbons-containing gaseous phase issuing from said first reaction zone through a bed of a Group VIII metal hydrogenation catalyst in a second reaction zone, at a temperature and a pressure in the same range as in the first zone, provided they are convenient to maintain said unreacted hydrogen and vaporized hydrocarbons in the gaseous state, thereby resulting in a highly pure gaseous cycloaliphatic hydrocarbon product substantially devoid of any aromatic hydrocarbon impurities, and

quantitative yield.

further resulting in a process requiring neither external heat nor large equipment normally associated with gas phase reactions.

2. The process of claim 1 wherein the temperature in the first reaction zone is from 10 to C. lower than the boiling temperature of the cycloliphatic hydrocarbon under the prevailing total pressure.

3. A process according to claim 1, wherein the molar ratio between the aromatic hydrocarbon and the corresponding cycloaliphatic hydrocarbon is kept substantially unchanged throughout the first reaction zone.

4. A process according to claim 1, wherein the temperature in the second zone is substantially higher than in the first zone.

5. A process according to claim 1, wherein the pressure in the second zone is substantially lower than in the first zone.

6. A process according to' claim 1, wherein the operating conditions of pressure and temperature are substantially the same in both reaction zones.

7. A process according to claim 1, wherein the aromatic hydrocarbon is benzene. A

8. A process according to claim 1, wherein the aromatic hydrocarbon is toluene.

9. A continuous process for the catalytic, non-destructive hydrogenation of an aromatic hydrocarbon, which process comprises the steps of:

(1) reacting hydrogen with an aromatic hydrocarbon selected from the group consisting of benzene, toluene, xylene, and naphthalene to produce the corresponding cycloaliphatic hydrocarbon, said reaction being conducted in a reaction medium comprising said aromatic hydrocarbon in the liquid phase, the correspbnding liquid cycloaliphatic hydrocarbon and a solid catalyst suspended in said liquids, said catalyst selected from the group consisting of nickel, cobalt, platinum, rhodium and ruthenium, the ratio of hydrogen to said aromatic hydrocarbon being at least stoichiometric; the molar ratio of said liquid aromatic hydrocarbon to said liquid cycloaliphatic hydrocarbon being substantially constant and being in the range of 0.0030.1 1, respectively, the operating temperature being 80250 C., the operating pressure being about 1-100 atmospheres, and with the pro- 3 8 vision that said operating temperature is l0100 C. lower than the boiling temperature of the cycloaliphatic hydrocarbon prevailing at the total reaction pressure;

(2) withdrawing from the reaction medium of step (1), a gaseous product stream comprising cycloaliphatic hydrocarbon, hydrogen, and unreacted aromatic hydrocarbon starting material;

(3) reacting said gaseous product stream in the gaseous phase in contact with a dry solid hydrogenating catalyst selected from the group consisting of nickel, cobalt, platinum, rhodium and ruthenium at about 80-250 C. and about 1-100 atmospheres with the provision that a gaseous phase be maintained, thereby resulting in a highly pure gaseous cycloaliphatic hydrocarbon product substantially devoid of any aromatic hydrocarbon impurities, and'further resulting in a process requiring neither external heat nor large equipment normally associated with gas phase reactions.

References Cited by the Examiner UNITED STATES iATENTS 2,952,625 9/60 Kelley et a1. 208-216 2,979,546 4/61 Grandio et al. 260667 3,054,833 9/62 Donaldson et a1 260-667 3,070,640 12/62 Pfeiifer et a1 260667 0 ALPHONSO D. SULLIVAN, Primary Examiner.

Claims

1. A PROCESS FOR THE CATALYTIC, NON-DESTRUCTIVE HYDROGENATION OF AT LEAST ONE AROMATIC HYDROCARBON, WHICH PROCESS COMPRISES THE STEPS OF: (1) INTRODUCING SAID AROMATIC HYDROCARBON IN THE LIQUID STATE INTO THE CORRESPONDING LIQUID CYCLOALIPHATIC HYDROCARBON CONTAINED IN A FIRST REACTION ZONE, PASSING THROUGH SAID CYCLOALIPHATIC HYDROCARBON, AN AMOUNT OF GASEOUS HYDROGEN AT LEAST EQUAL TO THE STOICHIOMETRIC RATIO WITH RESPECT TO THE AROMATIC HYDROCARBON, SAID LIQUID CYCLOALIPHATIC HYDROCARBON HAVING A GROUP VIII METAL HYDROGENATION CATALYST SUSPENDED THEREINTO, THE MOLAR RATIO BETWEEN SAID AROMATIC HYDROCARBON AND SAID CYCLOALIPHATIC HYDROCARBON IN SAID FIRST REACTION ZONE BEING MAINTAINED WITHIN THE RANGE OF FROM 0.003:1 TO 0.1:1, THE PRESSURE BEING COMPRISED BETWEEN ABOUT 1 AND 100 ATMOSPHERES AND THE TEMPERATURE COMPRISED BETWEEN 80 AND 250*C., PROVIDED THE TEMPERATURE IS AT LEAST 10*C. LOWER THAN THE INITIAL BOILING TEMPERATURE OF SAID CYCLOALIPHATIC HYDROCARBON UNDER THE PREVAILING PRESSURE, AND (2) PASSING AS SUCH THE WHOLE, UNREACTED HYDROGENAND VAPORIZED HYDROCARBONS-CONTAINING GASEOUS PHASE ISSUING FROM SAID FIRST REACTION ZONE THROUGH A BED OF A GROUP VIII METAL HYDROGENATION CATALYST IN A SECOND REACTION ZONE, AT A TEMPERATURE AND A PRESSURE IN THE SAME RANGE AS IN THE FIRST ZONE, PROVIDED THEY ARE CONVENIENT TO MAINTAIN SAID UNREACTED HYDROGEN AND VAPORIZED HYDROCARBONS IN THE GASEOUS STATE, THEREBY RESULTING IN A HIGHLY PURE GASEOUS CYCLOALIPHATIC HYDROCARBON PRODUCT SUBSTANTIALLY DEVOID OF ANY AROMATIC HYDROCARBON IMPURITIES, AND FURTHER RESULTING IN A PROCESS REQUIRING NEITHER EXTERNAL HEAT NOR LARGE EQUIPMENT NORMALLY ASSOCIATED WITH GAS PHASE REACTIONS.