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US2488083A - Manufacture of liquid hydrocarbons - Google Patents

Manufacture of liquid hydrocarbons

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US2488083A
US2488083A US67743446A US2488083A US 2488083 A US2488083 A US 2488083A US 67743446 A US67743446 A US 67743446A US 2488083 A US2488083 A US 2488083A
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chloride
line
methyl
hydrocarbons
per
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Manuel H Gorin
Gorin Everett
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ExxonMobil Oil Corp
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ExxonMobil Oil Corp
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    • CCHEMISTRY; METALLURGY
    • C07ORGANIC CHEMISTRY
    • C07CACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/26Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only halogen atoms as hetero-atoms
    • CCHEMISTRY; METALLURGY
    • C07ORGANIC CHEMISTRY
    • C07CACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2521/00Catalysts comprising the elements, oxides or hydroxides of magnesium, boron, aluminium, carbon, silicon, titanium, zirconium or hafnium
    • C07C2521/12Silica and alumina

Description

Patented Nov. 15, 1949 mal-s .Huil

MANUFACTURE OF LIQUID HYDROCARBONS Manuel H. Gorin and Everett Gorin, Dallas, Tex.,

assignors, by mesne assignments, to Socony- Vacuum Oil Company, Incorporated, New York, N. Y., a corporation of New York Application June 18, 1946, Serial No. 677,434

6 Claims.

This invention relates to the conversion of normally gaseous hydrocarbons to hydrocarbons of higher molecular Weight. More particularly. this invention relates to the conversion of methane and natural gas to a mixture of normally liquid hydrocarbons in a unitary process wherein the normally gaseous hydrocarbons are converted to the corresponding alkyl halides which are then condensed in a dehydrohalogeno-condensation reaction. The particular feature of the present invention is the recycle of at least a part of the normally gaseous primary condensation products for use as diluent of the alkyl chlorides in the condensation reaction. The term normally gaseous primary condensation products as used in the specification and claims refers to the C2, C3, and C4 hydrocarbons formed in the condensation process.

It is known to condense light alkyl halides by pyrolysis in strictly thermal conversions. For example, in the patent obtained by one of the present inventors, U. S. 2,320,274, entitled Conversion of normally gaseous hydrocarbons, it is taught to halogenate the hydrocarbons and then pyrolyze the light alkyl halides at temperatures above about 500 C. to produce benzene, acetylene, and ethylene. The method requires relatively high temperatures -in the pyrolysis step. Where it is desired to produce normally liquid hydrocarbons by the condensation of the alkyl halides such as methyl chloride and ethyl chloride lower temperatures are highly desirable.

It, therefore, is an object of our invention to provide an improved process for converting normally gaseous hydrocarbons such as methane and ethane to normally liquid hydrocarbons. Another object is to produce from methane and natural gas, hydrocarbons and hydrocarbon derivatives of higher molecular Weight, at least a part of which is utilizable in the form of recycle to the condensation reaction zone. A more specific object is to provide a process for condensing methane at temperatures below 500 C. to normally liquid hydrocarbons by halogenation of the methane, and condensing the methyl halide in the presence of recycled primary condensation products. These and other objects will be apparent from the description of our invention.

Methane is available in enormous quantities in natural gases. Natural gases are composed primarily of methane and ethane and usually contain progressively smaller amounts of propane, butane, pentane and higher hydrocarbons. Natural gases which contain the higher molecular weight hydrocarbons are usually termed Wet gases. After the removal of the condensible components, which may be used as such as motor fuel blending agents or which may be processed by dehydrogenation and/or isomerization to produce alkylation feed stock for the production of (Cl. 260-677l aviation blending naphthas, the residual noncondensible light hydrocarbons consisting of methane, ethane, and a small amount of propane are generally useful only as fuel. Hence, artificially dried natural gas as well as naturally occurring dry gas containing methane, ethane, and a small amount of propane may be used in our process.

We have found that the dimculties of prior art processes may be eliminated by employing a twostep operation wherein the light hydrocarbon such as methane or ethane or mixture of these is converted to a more reactive compound, and then this more reactive compound or mixture of compounds is catalytically treated to condense the hydrocarbon radicals of two or more of the more reactive intermediates. Thus, for example, we have found that if methane is rst converted to a methyl halide, the methyl halide may be catalytically condensed to hydrocarbons having two or more carbon atoms to the molecule. The corresponding hydrogen halide is eliminated from the methyl halide and may be reused for the production of methyl halide as described in the copending application of E. Gorin, entitled Hydrocarbon conversion process, Serial No. 507,618, led October 25, 1943, now U. S. Patent 2,407,828, and also described in the copending joint application of C. M. Fontana and E. Gorin, entitled Manufacture of halogenated hydrocarbons, Serial No. 548,351, led August "I, 1944. The first step of our process, namely, the halogenation of the paraliin hydrocarbons, may be carried out by other methods well known to those versed in the art.

The dehydrohalogeno-condensation of methyl chloride is effected by the aid of catalysts which possess both dehydrohalogenation and polymerization activity. Suitable catalysts are those which comprise an association of amphoteric and acidic oxides such as alumina on silica, zinc oxide on silica, boric oxide on alumina, beryllium oxide on silica or gallium oxide on silica. Other mixed oxide catalysts such as magnesia-silica, zirconiasilica, or thoria-silica may be used. Pure amphoteric or acidic oxides may be used alone, such as, i'Jr example, alumina, silica, or titania. Activecatalysts may also be prepared from naturally occurring materials such as by acid-treating clays of the fullers earth, bentonite, montmorillonite, or Attapulgus variety. Strictly acid-type catalysts such as phosphoric acid or pyrophosphoric acid, preferably mounted on a carrier such as kieselguhr, may also be used. However, we prefer to use the mixed oxide-type catalyst containing two or more oxides, particularly the activated alumina-silica catalyst.

The oxide catalysts are generally prepared in such a way that they are subjected to an acid environment in the last stages of their preparation in order that traces of alkali metal oxides which act ascatalyst poisons be substantially completely removed. However, when the only cation present during the preparation of the catalyst is the cation of a volatile base such as ammonia, this purication may be omitted. Alumina or alumina-silica catalysts containing more than traces of sodium are relatively inactive toward the dehydrohalogenation of methyl chloride. However, when the sodium is removed by treatment with dilute acid or by salts of weak bases such as ammonium chloride, aluminum chloride, or aluminum sulfate followed by water washing, drying, and calcining at about 500 C` these catalysts become highly active. Gel-type catalysts having high porosity are preferred, and it is highly desirable that the catalyst be capable of withstanding high temperatures such as those produced by oxidative regeneration, since considerable carbon is deposited on the catalyst in the reaction cycle which must be removed in order to restore the catalyst activity. Our preferred alumina gel and silica gel or combined alumina-silica catalysts possess the above-desired properties.

The alumina gel vis preferably prepared by forming an alumina sol. A mercury-aluminum amalgam is digested in a dilute acid such as acetic acid. The clear sol is separated from the unreacted amalgam and after setting to a gel the alumina gel is separated. Setting of the gel may be promoted by evaporation at 50 C. to 100 C. or by precipitation with dilute ammonia. Silica hydrogel may be prepared by addition of sodium silicate solution to an excess of mineral acid such as sulfuric acid. To prevent precipitation of the silica violent agitation and cooling of the acid may be used. These precautions may be omitted if acetic acid is used instead of mineral acid in which case the buffering action of the sodium acetate formed prevents precipitation of the silica. The moist washed gels of alumina and silica are ballmilled together, dried at 100 C., and activated by heating at about 500 C. for several hours to form the nished aluminasilica gel catalyst. The mol ratio of alumina to silica may be varied over a wide range since each of the pure gels taken separately are catalysts for the dehydrohalogeno-condensation reaction. It is generally preferred, however, to employ a considerable excess of silica gel over alumina gel and more particularly a mol ratio of about thirty to one to about five to one is preferred.

Other methods may be employed to combine the alumina and the silica. Thus, alumina may be deposited on the silica gel by hydrolytic precipitation of alumina from a dilute solution of an aluminum salt; that is, alumina may be precipitated on the silica gel by digestion of a sus pension of the gel in an aluminum salt solution. The alumina, on the other hand, may be precipitated on the silica by the addition of a basic reagent such as ammonia or ammonium carbonate to a suspension of silica gel in an aqueous solution of an aluminum salt. The mixed oxides of silica and alumina may be prepared as cogels by mixing solutions of sodium aluminate and sodium silicate and allowing the mixture to set to a clear hydrogel which is then broken to a relatively small particle size, thoroughly washed and treated with dilute acid to remove sodium therefrom, dried and activated by heating as described above.

The above catalysts are comparatively inactive at temperatures below 250 C. while in most cases temperatures above 300 C. are required to eiect a reasonably rapid reaction. Temperatures in excess of about 500 C. are unsuitable due to excessive decomposition of the products. Our preferred temperature range for operating the dehydrohalogeno-condensation reaction is from about 350 C. to about 425 C. A wide range of pressures may be employed in carrying out the conversion of alkyl chlorides to hydrocarbons having a greater number of carbon atoms, such pressures varying from one atmosphere to about forty atmospheres or higher.

The pressure employed in our process has a decided effect on the nature and distribution of the products. For example, when the process is carried out at high temperatures and low pressures with methyl chloride feed, the gaseous fraction consists primarily of low molecular weight olens such as ethylene, propylene, and butylene, and the gasoline fraction contains considerable unsaturates. When higher pressures, i. e., above six or seven atmospheres and temperatures below 400 C. are employed, the gaseous fraction contains relatively larger amounts of saturates than unsaturates, and the gasoline fraction is almost completely saturated. Under.

the latter conditions the lower boiling fraction,v

of the liquid product is highly branched and con-5 tains large amounts of such components as isopentane, 2,3-dimethyl butane and 2-methyl pentane and, hence, may be used as a blending stock for aviation gasoline. The heavier fractions are rich in aromatic hydrocarbons. Relatively large amounts of isobutane are formed under low temperature-high pressure operation while the lighter gases consist of a mixture of ethylene, propylene, ethane, and propane. In addition,

ethyl chloride is usually produced, the ratio of:

ethyl chloride to ethylene being higher in the case of low temperature-high pressure operation. We prefer to operate at pressures within the range of from about 15 to about 40 atmospheres. In general, only traces of alkyl chlorides boiling above ethyl chloride are produced when methyl chloride and ethyl chloride are condensed under the above conditions.

The optimum space velocity for carrying out. the dehydrohalogeno-condensation of alkyl halides depends on the particular alkyl halide, on the catalyst, and on the temperatureand pres-- sure employed. For methyl chloride the space. velocity expressed in terms of volumes of methyll chloride gas at standard temperature and pressure passed per volume of catalyst per minute: is in general higher the higher the temperature;

and the higher the pressure, and for our preferred alumina-silica gel catalyst lies within the range of from about 5 to about 50 volumes of gaseous methyl chloride per volume of catalyst space per minute at temperatures within the range of 350 C. to 425 C. and pressures from one atmosphere to twenty atmospheres. When operating at pressures of from about twenty atmospheres to about forty atmospheres within the temperature range of 370-400 C. with the alumina gel-silica gel catalyst, we prefer space velocities within the.

range of from about 5 to 20 volumes of gaseous methyl chloride per volume of catalyst space per minute.

The light oleflns such as ethylene, propylene,`4 and butylene produced in the reaction may be'- cycle the isobutane and ethyl chloride produced ommen fiUUaVI ananas since these materials are likewise converted to liquid: hydrocarbons by condensation with the methyl chloride or with primary products such as ethylene formed from the methyl chloride. We have found that higher yields of the liquid hydrocarbons are obtained when the condensation of methyl chloride is carried out in the presence of active diluents such as ethylene, propylene, isobutane, and ethyl chloride which are reacted to form hydrocarbons of four or more carbon atoms. Inactive diluents such as propane and normal butane likewise have a favorable effect on the reaction due to their dilution effect resulting in keeping the instantaneous concentration of methyl chloride at a low figure which is desirable for the most successful operation of our process. We also maintain a low instantaneous concentration of methyl chloride by introducing the vaporized feed at a multiplicity of points to the stationary catalyst bed described in our preferred method of operation below.

Referring now to the drawing, towers I and 2 contain alumina gel-silica gel catalyst either as a continuous bed or disposed in trays. We prefer the latter method of packing these towers, since the difficulty of controlling regeneration temperatures is reduced if the catalyst is maintained in a series of relatively shallow beds. Towers I and 2 are manifolded in such a manner that while one of these is on stream for the condensation reaction the other is undergoing oxidation regeneration thus providing for substantially continuous operation. The hydrocarbon halogenation zone is represented diagrammatically by 3. As indicated hereinabove, dry natural gas containing methane and ethane is introduced to zone 3 through line 5. Chlorine or hydrogen chloride and oxygen are introduced to the chlorination zone through line 6. Zone 4 is provided for the separation of hydrogen chloride from Cz hydrocarbons in the product.

Liquid methyl chloride is passed from zone 3 through valved line I and line II, and is discharged at a pressure of from about 500 to about 600 pounds per square inch by means of pump I2, whence it is passed through heat exchangers I3 and I4 for absorbing heat from the reaction product in line I5 and hot regeneration gases in line I6. From exchanger I4 the hot feed product passes via line I1 to manifold line I8 provided with valves I9 and 20 and connected with reactor feed manifold lines 2| and 22. With valve I EI in line I8 open and valve 20 closed, the hot methyl chloride vapor is introduced t0 tower I by a multiplicity of valved feed lines leading from manifold line 2| at a temperature Within the range of 370 C.400 C. and a space velocity within the range of 5 to 20 volumes of methyl chloride per volume of catalyst space per minute.

We have found that both inactive and reactive diluents aid in the condensation of methyl chloride and hence, we introduce isobutane and ethyl chloride product from the primary condensation through line 30 which joins manifold line 3I provided with valve 32, which is open for transfer of diluent, and valve 33 which is closed. Where it is desirable to eliminate excess propane and normal butane from the process, isobutane diluent from the product may be furnished in the form of methyl chloride-isobutane azeotrope, as described hereinbelow, via recycle line I03 which connects with feed line II. The mol ratio of hydrocarbon diluent recycle stock to total alkyl chloride fed to reactor I may be from about 0.5 to 1 to about 3 to 1 while the internal ratio of diluent to alkyl chloride is preferably maintained within the range of from about 4 to 1 to about 10 to 1 or higher. Under this mode of operation from to about 90 per cent of the alkyl chloride is decomposed per pass and from to 85 per cent of the alkyl chloride decomposed forms hydrocarbons having a greater number of carbon atoms than the alkyl chlorides in the feed to the reactor, 15 to 20 per cent appearing as carbon and methane decomposition products. The use of diluents, either active or inactive in the condensation reaction, reduces greatly the amount of carbon and methane formation, thereby directing the condensation reaction toward the formation of the more valuable liquid hydrocarbon product.

In our process the diluents are supplied exclusively from the normally gaseous hydrocarbons produced in the product, which hydrocarbons are characterized by having more than one carbon atom and less than five carbon atoms per mole` cule. We also recycle unreacted alkyl chloride reactants. The inactive diluents such as ethane. propane, and normal butane are usually recycled only in part. Since the separation of a mixture of ethylene and ethane from hydrogen chloride is difficult and since the mixture of these gaseous components of the product may be used advantageously to produce alkyl chloride feed to the condensation reaction, the ethane and ethylene are recycled only in part to the condensation reaction zone. Thus, We prefer to recycle a part of the ethylene and substantially all of the propylene and isobutane, and to withdraw from the recycle stream excess propane and normal butane which tend to build up in the system. At least a part of these parafiinic hydrocarbons may be returned to the reaction zone to maintain in the feed thereto a ratio of hydrocarbons to alkyl chlorides within the range of from 0.5 to 1.0 to 3.0 to 1.0.

The product from reactor I consisting of a mixture of methane, hydrogen chloride, ethylene, ethane, ethyl chloride, propane, propylene, unreacted methyl chloride, isobutane and other C4 hydrocarbons along with higher boiling hydrocarbons is passed by drawoi line 40 to manifold product line 4I which is equipped with valves 42 and 43, the latter being closed and the former open for delivery of the product to line I5 leading to heat exchanger I3 where the product is subjected to initial cooling. From exchanger I3 the product passes via line 44 to cooler 45 and thence by line 46 to fractionator 41. Fractionator 41 which is provided with indirect heating means 48 operates as a stripping column for the removal of light gases from the reaction product, the pressure maintained therein being somewhat lower than the pressure in reactor I. These light gases comprising methane, ethylene, ethane, and a part of the hydrogen chloride product of the reaction pass overhead through line 49 Jo condenser 50 and thence by line 5I to reflux drum 52. Overhead gas consisting primarily of methane is recycled through valved line 53 to line 5 and thence to the methane chlorination process for the production of methyl chloride feed to the process. Condensate ln reflux drum 52 is removed there- -from by pump 54 in line 55 to be returned in part to tower 41 as reflux through valved line 56, a part of excess condensate being sent via valved line 51 and line 6 to the methane chlorination process described in the aforementioned copending application, Serial No. 548,351, or to other hydrogen chloride recovery systems. If desired, a part of the stream in line 51 which contains the C2 hydr'oarbons and hydrogen chloride may be passed to zone 4 through line 58 Where hydrogen chloride is separated from these gases by any means well known in the art, such as, by absorption. The hydrogen chloride is recycled to chlorination zone 3 via lines 59, 51, and 6, and the C.: hydrocarbons are recycled to the condensation zone as diluent through lines 59a, 30, and 3|. It has been found that large amounts of hydrogen chloride have an inhibiting effect on the course of the condensation reaction, and it is usually desirable to maintain the hydrogen chloride concentration below 25 per cent of the total gas and in any case hydrogen chloride should not constitute over 50 mol per cent of the total gas eiiluent from reactor I. Hence, hydrogen chloride is removed as far as practicable from recycle stock in order to hold the concentration of hydrogen chloride in reactor to a minimum.

The liquid product from fractionator 41 is transferred via pressure release valve 60 in line 6| to butanizer tower 62 for the removal of propane, propylene, the C4 fraction, unconverted methyl chloride and ethyl chloride from the higher boiling product. These products are all suitable for recycle, at least in part, to reactor I either as diluents to promote the reaction or as condensable alkyl chlorides. The overhead from tower 62 will also contain hydrogen chloride in minor amounts which is retained in the bottom product from tower 41 as a result of the relatively high pressure fractionation in tower 41. Tower 62 is operated at a pressure of about 200 pounds per square inch and is provided with reboiler means 63. The overhead gaseous product from tower 62 passes via line 64, condenser 65, and line 66 to reflux drum 61 from which noncondensed gas, primarily hydrogen chloride passes via valved line 68 to line 51 for use in producing methyl chloride as described in the afore-mentioned copending applications. condensate from reflux drum 61 is picked up by pump 69 in line 10, a part being transferred through valved line 1| to tower 62 as reux and the remainder passed via valved line 12 to recycle line 30. As stated hereinabove, the components of the stream in line 12 are all suitable for recycle to the reaction zone. Hence, valved recycle line 30 is provided, which leads from line 12 directly to manifold line 3|. If desired, this recycle product may be diverted to the main feed stream in line I I via valved line 29 leading from line 30. The C3 and C4 hydrocarbons, ethyl chloride, and, if desired, a part of the Cz hydrocarbons are recycled as described above until the concentration of diluent in condensation tower I is built up to the hereinabove described ratio of diluent to virgin and recycle alkyl chloride. Excess diluent such as propane and normal butane produced in the process are eliminated therefrom as described hereinbelow.

When the propane and/or normal butane content of the stream in line 12 becomes excessive, a part of said recycle stream is diverted to tower 13 to separate propane and propylene from the C4 hydrocarbons and alkyl chlorides. Tower 13 is maintained at a pressure of from about 200 to 250 pounds per square inch with tower top temperature maintained at about 35 C. to 40 C., tower 13 being provided with reboiler means 14. The overhead consisting of propylene and propane passes via line 15, cooler 16 and line 11 to reflux drum 18. Uncondensed propylene passes via line 19 to recycle line 30 and the condensate is removed from drum 18 by means of pump 9| in line 82 for transfer in part to tower 13 as reflux via valved line 83 and the remainder, comprising relatively pure propane, may be sent to storage via line 84. When it is desired to increase the ratio of propane diluent to normal butane in reactor I a part of propane in line 84 is introduced to line 3U and recycled to the condensation reaction.

The bottom product from tower 13 is transferred via line and pressure release valve 9| to deisobutanizer tower 92 which is provided with reboiler means 93 and which is operated at pressures within the range of 100 to 150 pounds per square inch. We have found that methyl chloride forms an azeotrope with isobutane and hence the overhead from tower 92 consists primarily of methyl chloride-isobutane azeotrope boiling at about -25.6 at atmospheric pressure, provided these components are present in the proper proportions. This azeotrope consists of approximately '77 mol per cent methyl chloride and 23 mol per cent isobutane. Provision is made for adding methyl chloride to line 90 via valved leadoff line 94 from line |I by means of pump 95 for augmentation of methyl chloride to tower 92 when the conversion in reactor I is suiciently high that there is insufficient methyl chloride to form the desired azeotrope with isobutane. Where it is undesirable to recycle all of the isobutane the excess isobutane not combined as the azeotrope may be removed along with normal butane through line |05 described below. The separation of isobutane from the other C4 hydrocarbons and ethyl chloride is sharp. Hence, the overhead product from tower 92 will contain only a trace of propane and the azeotrope plus any of the uncombined components of the azeotrope mixture, that is, free isobutane or free methyl chloride. The overhead passes through line 96, condenser 91, and line 98 to reflux drum 99 whence the condensate is sent in part by means of pump |00 in line |0| and via valved line |02 to tower 92 as reflux, the remainder being passed via valved line |03 for recycle through line |I to reactor I as described hereinabove. Any accumulation of gaseous propane in drum 99 is removed via line |04 to storage. The C4 bottom fraction which is removed from tower 92 via line |05 contains some ethyl chloride which may be separated by scrubbing with an alcohol-water mixture and the separated ethyl chloride may be recycled to reactor I while the C4 fraction may be subjected to alkylation and/or isomerization reactions to produce aviation grade gasoline. If desired, a part of the C4 bottom fraction from tower 92 may be recycled to reactor in order to maintain the desired ratio of hydrocarbon to alkyl chloride in the feed as described hereinabove.

Returning now to debutanizer tower 62, the bottom product consisting of C5 and higher boiling hydrocarbons is withdrawn via line ||0 leading to wash tower where it is washed countercurrently with an aqueous caustic solution introduced to tower through line ||2 to remove any traces of hydrogen chloride. The C5|naphtha passes overhead to storage through line ||3 and the partially spent caustic wash is withdrawn vla valved line I I4 for suitable disposal.

After reactor has been on stream for an interval up to 90 minutes depending on the operating conditions relative to space velocity and pressure, sulcient carbon will have accumulatstabili-i R005;

ed on the catalyst; to necessitate reactivation. This is accomplished by passing air through the reactor, .the air being diluted with nue gas in order t6 rcontrol the rate of oxidation and to maintain the regeneration temperature below about 650 C., preferably below about 600 C. However, reactor I and the catalyst therein will contain some hydrogen chloride and other reaction products which may be recovered by purging the reactor prior to oxidative reactivation. It is desirable to remove the hydrogen chloride and the residual hydrocarbons prior to burning the carbon from the catalyst since the combustion of the hydrocarbons results in the formapounds per square inch and preferably at pressures of about 500 pounds per square inch, we do not wish to be limited to this type operation. The dehydrohalogeno-condensation reaction may be carried out at pressures below 300 pounds, for example, of the order of atmospheric to 50 pounds per square inch. However, such type operation involves greater expenditureof capital for compression equipment for separation of predominantly gaseous hydrocarbon product when operating at low pressures and higher temperatures. Certain advantages favor low pressure operation. For example, the carbon deposited on the catalyst may be removed more advantageoustion of water vapor and the presence of water 1y by operating the oxidative reactivation step vapor with hydrogen chloride aggravates corroby the Well-known hindered flow-type DIOCedure Sion problems, which is usually adaptable at pressures below '75 Tower I is isolated for purging before regeneror 100 pounds per square inch. In this type of ation by opening valve in line |28, and clos- Operation the catalyst is in the form of powder ing valve |24 in line |22, valve 32 in line 3|, 20 which is introduced to the regeneration zone susvalve |44 in line |42, valve |41 in line |46, valve 42 in line 4|, and valve |29 in line |27. Methane purge gas is introduced via line |2| and passes via lines |22 and 3| to reactor for removal of residual hydrogen chloride and hydrocarbon vapors. The purge gas leaves reactor through line and thence passes via line |28 to line 5 for transfer to the methyl chloride production zone 4. The methane purging operation is followed by a short flue gas purge in order to clear reactor of methane prior to the reactivation step.

-Referring now to the reactivation operation which is carried out in tower 2 while reactor is on stream for the condensation reaction and for thepurging operation, a mixture of air and ue gas is introduced to the system by means of compressor |40 in line 14|. The extent of dilution of oxygen in the regenerating gas should be such that temperatures in excess of 650 C. are avoided and preferably the temperature should not exceed 600 C. since synthetic alumina-silica catalyst and also acid-treated clays tend to become permanently deactivated at higher temperatures. Control of reactivation temperatures is more readily obtained by disposing the catalyst in a series of relatively shallow beds rather than as a. continuous bed. The ue gas diluted air is passed from line |4| to manifold line |42, which is provided with valves |43 and |44, and thence to line 22 for introduction to tower 2 via manifoldland valved feed lines leading from line 22. The regeneration zone is isolated by closing valve |44 in line |42, valve 20 in line I8, valve |41 in line |46, valve 43 in line 41, and valve |48 in line |21, and provision for continuous flow of regeneration gas is made by maintaining valve |43 in line |42 and valve 145 in line |46 in the open position. The regeneration gas passes from tower 2 pended in the regeneration gas and maintained therein by regulating the flow of gas through the regeneration zone at such a velocity that the fiuidized powder assumes physical properties not unlike those of a boiling liquid. Oxidative regeneration temperatures are more uniform and more easily controlled in such type operation. This type operation may also be applied to the dehydrohalogeno-condensation side of the process if the pressure is maintained below 75 or 100 pounds per square inch.

We may also adopt other methods of catalytic contacting to our process of converting methyl chloride to normally liquid hydrocarbons. For example, the various types of moving catalyst bed techniques well known in the art of hydrocarbon conversion, particularly in the art of converting petroleum hydrocarbons and hydrocarbon mixtures, may be used if relatively low pressure operation is followed.

The following examples are introduced to show the effects of pressure and diluents in carrying out our process. Separately prepared moist silica gel and alumina gel were ball-milled in a mol ratio of silica gel to alumina gel of approximately 8 to 1, the mixture was dried at a temperature below 100 C. and finally activated by calcining at 500 C. for about 20 hours. 'This catalyst was used in obtaining the following results.

Example 1 Methyl chloride was passed over a sample of the above alumina-silica gel catalyst at 425 C. and at a space velocity of 4.3 volumes of methyl chloride vapor per volume of catalyst space per minute, the reaction zone being maintained at atmospheric pressure. 24.5 per cent of the methyl chloride was decomposed to give a product which analyzed as follows:

via' line |49 and is passed via lines |46 and I6 to 60 heat exchanger I4, whence it is eliminated from Product @aglagggggoosfls the system through valved line |50. If desired,

a part of the spent regeneration gas may be re- 191 cycled to line |4| through valved line I5| for as dilution of fresh regeneration air. The ilow of l? regeneration gas is continued for a temperature 5.1 adjustment period following complete reactivation of the catalyst until the catalyst bed tem- 3E perature has lowered to such a level that fresh m0 feed will be held within the desired temperature range for the condensation reaction. The purge gas should contain no oxygen during this purging operation.

Although we have described our invention as being carried out at pressures in excess of 300 l 1 were accounted for in hydrocarbon product boiling above the C4 fraction, that is, in the normally lig-uid hydrocarbon fraction.

Example 2 Methyl chloride was passed over a sample of the above alumina-silica gel catalyst at 300 pounds per square inch pressure at a space velocity of approximately 8 volumes of methyl chloride vapor per volume of catalyst space per minute, the temperature being maintained at 370 C. 47.5 per cent of the methyl chloride was decomposed. The product was analyzed as follows:

12 molecular weight normally liquid hydrocarbons are formed.

Example 4 A mixture of 51.1 mol per cent methyl chloride and 48.9 mol per cent of ethylene was passed over the alumina-silica gel catalyst at atmospheric pressure and 425 C. at a space velocity of 4.4 volumes of methyl chloride vapor plus ethylene per volume of catalyst space per minute. 44.4 per cent of the methyl chloride was decomposed to give the following yields:

Mols Carbon/100 mols Product CHaCl decomposed The C4+ fraction was completely saturated which illustrates the effect of increasing the pressure on the degree of saturation of the product. The total yield of C4+ product amounted to 40.4 per cent of the theoretical as compared with the yield of 28.8 per cent on a corresponding basis in Example 1 which was carried out at atmospheric pressure. Hydrocarbon product boiling above the C4 fraction accounted for 17.5 mols of carbon per mol of 100 mols of methyl chloride decomposed. A greater proportion of the carbon of the methyl chloride feed appears in the C4 fraction of the product than in Example 1 which C4 fraction consists essentially of isobutane.

Example 3 A mixture containing 69 mol per cent of propane and 31 mol per cent of methyl chloride was passed over the above described alumina-silica gel actalyst at a space velocity of about 4.6 volumes of methyl chloride vapor plus propane per volume of catalyst space per minute and at a temperature of 425 C., atmospheric pressure being maintained in the reactor. 61.2 per cent of the methyl chloride was decomposed. The yields of products are given below:

Mols Carbon/100 mols Product CHlCl decomposed Example 5 A mixture of 65.7 mol per cent of isobutane and 34.3 mol per cent of methyl chloride was passed over the above alumina-silica gel catalyst at 401 C., 300 pounds per square inch pressure and at a space velocity of 32.4 volumes of methyl chloride plus isobutane per volume of catalyst space per minute. 51.2 per cent of the methyl chloride was decomposed to yield the following products:

Mols Carbon/ mols Product CHaCl decomposed Approximately 12.6 mols of isobutane were consumed per 100 mols of methyl chloride decomposed. In addition, a small fraction of the isqbutane was isomerized to normal butane. The gasoline fraction was completely saturated and consisted primarily of a mixture of branched chain paramns. However, the C1+ fraction was rich in aromatlcs.

Comparing Examples 1 and 2 it is indicated that operation under pressure favors the production of paraflins rather than olens. Thus, it can be seen that the recycle of ethylene, propane, and the C4 fraction promotes the Iormaltion of normally liquid hydrocarbons and the production of carbon ls greatly decreased by the presence of such diluents.

This Tapplication is a continuation-in-part of our copending application entitled Catalytic conversion of normally gaseous hydrocarbons, Serial No. 556,746, led October 2, 1944, now abandoned.

We claim:

1. The process for the conversion of methyl chloride to a mixture of normally liquid hydrocarbons which comprises the steps of (l) passing a feed stream consisting essentially of methyl chloride in contact with an alumina-silica catalyst in a reaction zone at a temperature within the range of from about 250 C. to about 500 C. to form a mixture comprising methane, normally gaseous parafns higher than methane, normally gaseous olens, unreacted methyl chloride, ethyl chloride and normally liquid hydrocarbons, (2) fractionating the product of step l to obtain a stream containing C3 hydrocarbons, C4 hydrocarbons, methyl chloride and ethyl chloride and a stream consisting essentially of normally liquid hydrocarbons, (3) fractionating the rst mentioned stream f step (2) to separate therefrom propylene, ethyl chloride, and a methyl chloride-isobutane azeotrope, (4) recycling the propylene, ethyl chloride and methyl chlorideisobutane azeotrope from step 3 to step 1, and recovering the normally liquid hydrocarbons from step 2.

2. The process for the conversion of an alkyl chloride feed stock consisting essentially of alkyl chlorides having less than four carbon atoms per molecule and consisting at least predominantly of methyl chloride which comprises the steps of (1) passing a feed stream consisting of said alkyl chloride feed stock through a reaction zone at a temperature Within the range of from 250 C. to 500 C. in contact With a catalyst consisting of an association of silica with an oxide of a metal selected from the class consisting of aluminum, zinc, beryllium, gallium, magnesium, zirconium and thorium to form a product mixture comprising methane, normally gaseous normal parans, isobutane, normally gaseous olens, unreacted alkyl chlorides, hydrogen chloride, and normally liquid hydrocarbons boiling in the gasoline range, (2) recycling at least a part of the olens and the isobutane in the product eflluent to the reaction zone, 3) recycling at least; a part of the unreacted alkyl chlorides to the reaction zone, (4) recycling a sufficient part of the normally gaseous normal parafflns in the product eilluent to maintain in the total feed to the reaction zone a mol ratio of hydrocarbons to total alkyl chlorides within the range of from about 0.5 to 1.0 to about 3.0 to 1.0 and (5) recovering the normally liquid hydrocarbons from the product eluent from step l.

3. The process for tde conversion of an alkyl chloride feed stock consisting essentially of alkyl chlorides having less than four carbon atoms per molecule and consisting at least predominantly of methyl chloride which comprises the steps of (l) passing a feed stream consisting of said alkyl chloride feed stock through a reaction zone at a pressure of from to 40 atmospheres and at a temperature within the range of from 250 C. to 500 C. in contact with a catalyst consisting of an association of silica with an oxide of a metal selected from the class consisting of aluminum, zinc, beryllium, gallium, magnesium, zirconium and thorium to form a product mixture comprising methane, normally gaseous normal paraflns, isobutane, normally gaseous olens, unreacted alkyl chlorides, hydrogen chloride, and normally liquid hydrocarbons boiling in the gasoline range, (2) recycling at least a part of the olefins and the isobutane in the product eliluent to the reaction zone, (3) recycling at least a part of the unreacted alkyl chlorides to the reaction zone, (4) recycling a sufcient part of the normally gaseous normal parafns in the product eilluent to maintain in the total feed to the reaction zone a mol ratio of hydrocarbons to total alkyl chlorides within the range of from about 0.5 to 1.0 to about 3.0 to 1.0 and (5) recovering the normally liquid hydrocarbons from the product effluent from step 1.

4, The process for the conversion of methyl chloride to a mixture of normally liquid hydrocarbons which comprises the steps of (1) passing a feed stream consisting essentially of methyl chloride in contact with an alumina-silica catalyst in a reaction zone at a temperature within the range of from about 250 C. to about 500 C. and at a pressure of from about 15 to 40 atmospheres to form a mixture comprising methane, normally gaseous normal parafns, isobutane, normally gaseous olens, unreacted methyl chloride. ethyl chloride, and normally liquid hydrocarbons boiling in the gasoline range, (2) recycling at least a part of the olens and the isobutane in the product eflluent to the reaction zone, (3) recycling at least a part of the methyl chloride and ethyl chloride to the reaction zone, (4) recycling a suicient part of the normally gaseous normal parafns in the product edluent to maintain in the total feed to the reaction zone a mol of hydrocarbons to total alkyl chlorides Within the range of from about 0.5 to 1.0 to about 3.0 to 1.0 and (5) recovering the normally liquid hydrocarbons from the product effluent of step 1.

5. The process of claim 2 in which the catalyst is alumina-silica.

6. The process of claim 3 in which the catalyst is alumina-silica.

MANUEL H. GORIN. EVERETT GORIN.

REFERENCES CITED The following references are ol' record in the le of this patent:

UNITED STATES PATENTS Number Name Date 1,879,912 Schmidt et al Sept. 27, 1932 2,255,834 Carmody et al Sept. 16, 1941 2,364,762 Schmerling et al. Dec. 12, 1944 2,417,119 Miller et al Mar. 11, 1947 OTHER REFERENCES Lu et al., Ref. & Nat. Gas. Mfgr. 20, No. 9, pages 13G-133.

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US2658090A (en) * 1950-02-15 1953-11-03 Ruhrchemie Ag Production of olefins by dechlorination of alkyl chlorides
US4199533A (en) * 1978-11-03 1980-04-22 University Of Southern California Conversion of methane
WO1985002399A1 (en) * 1983-11-30 1985-06-06 The British Petroleum Company P.L.C. Process for the production of hydrocarbons from c1 to c4 monohaloalkanes
US4714796A (en) * 1987-04-21 1987-12-22 Illinois Institute Of Technology Production of higher molecular weight hydrocarbons from methane
US4769504A (en) * 1987-03-04 1988-09-06 The United States Of America As Represented By The United States Department Of Energy Process for converting light alkanes to higher hydrocarbons
US4795843A (en) * 1985-08-26 1989-01-03 Uop Inc. Conversion of methane into larger organic hydrocarbons
US4804797A (en) * 1987-08-24 1989-02-14 Gas Research Institute Production of commodity chemicals from natural gas by methane chlorination
US4973786A (en) * 1987-10-19 1990-11-27 Karra Sankaram B Process for the pyrolytic oxidation of methane to higher molecular weight hydrocarbons and synthesis gas
US4983783A (en) * 1989-07-25 1991-01-08 Illinois Institute Of Technology Reduction in carbon oxides in oxidative pyrolysis of halogenated methanes
US5068478A (en) * 1990-05-25 1991-11-26 Energia Andina, Ltd. Producing alkenes and alkynes from alkanes and alkenes
US5157189A (en) * 1987-10-19 1992-10-20 Karra Sankaram B Conversion of light hydrocarbons to higher hydrocarbons
US5663472A (en) * 1985-07-22 1997-09-02 University Of Southern California Production of alkenes
US5705728A (en) * 1990-12-06 1998-01-06 Occidental Chemical Corporation Process for the production of ethylene and mixture containing ethylene
US20050234276A1 (en) * 2004-04-16 2005-10-20 Waycuilis John J Process for converting gaseous alkanes to liquid hydrocarbons
US20060100469A1 (en) * 2004-04-16 2006-05-11 Waycuilis John J Process for converting gaseous alkanes to olefins and liquid hydrocarbons
US20070100189A1 (en) * 2005-10-27 2007-05-03 Stauffer John E Methyl bromide to olefins
US20090069606A1 (en) * 2005-04-11 2009-03-12 Grt, Inc. Method of making alkoxylates
US7579510B2 (en) 2006-02-03 2009-08-25 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US20090247796A1 (en) * 2004-04-16 2009-10-01 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons
US20090308759A1 (en) * 2008-06-13 2009-12-17 Marathon Gtf Technology, Ltd. Bromine-based method and system for converting gaseous alkanes to liquid hydrocarbons using electrolysis for bromine recovery
US20090312586A1 (en) * 2008-06-13 2009-12-17 Marathon Gtf Technology, Ltd. Hydrogenation of multi-brominated alkanes
US7674941B2 (en) 2004-04-16 2010-03-09 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons
US7838708B2 (en) 2001-06-20 2010-11-23 Grt, Inc. Hydrocarbon conversion process improvements
US7847139B2 (en) 2003-07-15 2010-12-07 Grt, Inc. Hydrocarbon synthesis
US20110015458A1 (en) * 2009-07-15 2011-01-20 Marathon Gtf Technology, Ltd. Conversion of hydrogen bromide to elemental bromine
US7883568B2 (en) 2006-02-03 2011-02-08 Grt, Inc. Separation of light gases from halogens
US7964764B2 (en) 2003-07-15 2011-06-21 Grt, Inc. Hydrocarbon synthesis
US7998438B2 (en) 2007-05-24 2011-08-16 Grt, Inc. Zone reactor incorporating reversible hydrogen halide capture and release
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US8273929B2 (en) 2008-07-18 2012-09-25 Grt, Inc. Continuous process for converting natural gas to liquid hydrocarbons
US20120313034A1 (en) * 2011-06-10 2012-12-13 Marathon Gtf Technology, Ltd. Processes and Systems for Demethanization of Brominated Hydrocarbons
US8367884B2 (en) 2010-03-02 2013-02-05 Marathon Gtf Technology, Ltd. Processes and systems for the staged synthesis of alkyl bromides
US8642822B2 (en) 2004-04-16 2014-02-04 Marathon Gtf Technology, Ltd. Processes for converting gaseous alkanes to liquid hydrocarbons using microchannel reactor
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US2255834A (en) * 1938-11-05 1941-09-16 North American Rayon Corp Method of rubberizing cellulose fabrics
US2364762A (en) * 1941-05-19 1944-12-12 Universal Oil Prod Co Treatment of aromatic hydrocarbons
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US2658090A (en) * 1950-02-15 1953-11-03 Ruhrchemie Ag Production of olefins by dechlorination of alkyl chlorides
US4199533A (en) * 1978-11-03 1980-04-22 University Of Southern California Conversion of methane
WO1985002399A1 (en) * 1983-11-30 1985-06-06 The British Petroleum Company P.L.C. Process for the production of hydrocarbons from c1 to c4 monohaloalkanes
EP0144219A1 (en) * 1983-11-30 1985-06-12 The British Petroleum Company p.l.c. Process for the production of hydrocarbons from C1 to C4 monohaloalkanes
US4579996A (en) * 1983-11-30 1986-04-01 The British Petroleum Company P.L.C. Process for the production of hydrocarbons from C1 to C4 monohaloalkanes
US5663472A (en) * 1985-07-22 1997-09-02 University Of Southern California Production of alkenes
US4795843A (en) * 1985-08-26 1989-01-03 Uop Inc. Conversion of methane into larger organic hydrocarbons
US4769504A (en) * 1987-03-04 1988-09-06 The United States Of America As Represented By The United States Department Of Energy Process for converting light alkanes to higher hydrocarbons
US4714796A (en) * 1987-04-21 1987-12-22 Illinois Institute Of Technology Production of higher molecular weight hydrocarbons from methane
US4804797A (en) * 1987-08-24 1989-02-14 Gas Research Institute Production of commodity chemicals from natural gas by methane chlorination
US4973786A (en) * 1987-10-19 1990-11-27 Karra Sankaram B Process for the pyrolytic oxidation of methane to higher molecular weight hydrocarbons and synthesis gas
US5157189A (en) * 1987-10-19 1992-10-20 Karra Sankaram B Conversion of light hydrocarbons to higher hydrocarbons
US4983783A (en) * 1989-07-25 1991-01-08 Illinois Institute Of Technology Reduction in carbon oxides in oxidative pyrolysis of halogenated methanes
US5068478A (en) * 1990-05-25 1991-11-26 Energia Andina, Ltd. Producing alkenes and alkynes from alkanes and alkenes
US5705728A (en) * 1990-12-06 1998-01-06 Occidental Chemical Corporation Process for the production of ethylene and mixture containing ethylene
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US7838708B2 (en) 2001-06-20 2010-11-23 Grt, Inc. Hydrocarbon conversion process improvements
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US20090069606A1 (en) * 2005-04-11 2009-03-12 Grt, Inc. Method of making alkoxylates
US7683230B2 (en) * 2005-10-27 2010-03-23 Stauffer John E Methyl bromide to olefins
US20070100189A1 (en) * 2005-10-27 2007-05-03 Stauffer John E Methyl bromide to olefins
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