US2864875A

US2864875A - Hydroisomerization

Info

Publication number: US2864875A
Application number: US277304A
Authority: US
Inventors: Joseph B Mckinley; William A Horne
Original assignee: Gulf Research and Development Co
Current assignee: Gulf Research and Development Co
Priority date: 1952-03-18
Filing date: 1952-03-18
Publication date: 1958-12-16
Anticipated expiration: 1975-12-16

Description

Dec. 16, 1958 J. B. MCKINLEY ET AL 2,864,875

HYDROISOMERIZATION Filed March 18. 1952 nited States HY DROISOMERIZATION Application March 18, 1952, Serial No. 277,394

1 Claim; (CLIM-683.65)

This invention relates to a catalytic hydroisomerization processfor the treatment of-hydrocarbons to convert them to more-valuable products. More particularly, the invention relates to a process forl the `catalytic hydroisomerization ofhydrocarbons of the naphtha or gasoline boilingrange `to increase the value of such hydrocarbons as components of fuelsfor internal combustion engines.

The processlof the invention can be employed for the isomerizationof a singleV straight chain or slightly branched chain paraflinor a mixture of such paraiiins and alsov to -theisomeriZation-of-.a single naphthene or a mixture of naphthenes. It can also be employed for the conversion of 'natural or casing-head gasolines, which are'4 usually largely paranic in composition, and for the conversion of straight run petroleum naphthas. Naphtha or gasoline fractions which contain a relatively largey proportion of parafns, for example 50 percent or more, are especially valuable charge stocks for the present process. However, the advantages of the invention are "obtained at least to some extent when charging other naphthas. Regardless ofthe charge stock to the process, the process is characterized'in that high yields of good quality productsare obtained.

A-swill appearfrom Ythe following-description, when a naphtha' fraction is-employed as the charge. stock, the presentprocess is related to hydroreforming` or hydroforming. 1n'hydroformiug, however, the principal objective has been to promote reactions yielding aromatics, such asithe dehydrogenat-iou of naphthenes and the dehydrocyclizat-ionof parans, at the expense of isomerization reactionsand 'with' necessary concomitant cracking of Vcomponents of the charge stock, particularly the paraiiinic constituents. In the--present process, conditions are maintained, includingvhigh temperatures, which are conducive to the isomerizationfof paraiiins andnaphthenes while-avoiding excessive cracking andl dehydrogenation of the'rcharge-stock.v Thus, while dehydrogenation reactions take place in the present process, the dehydrogenationo'a naphthene constituent of the charge, such as methyl'cyclopentane; will o-ccur after ysubstantial isomeriz'ation has takenl place.` As a result, a substantial proportion of methyl cyclope-ntane is converted to benzene;v Simple'fdehydrogenation'reactions takelplace principally with respect to the six carbon atom naphthene ring compounds, yielding'hydrocarbonsof the'benzene series.

The Yprocess ofthe' invention comprises continuously ilowing into contact with a metal oxide catalyst of the classdescribedhereinafter, hydrogeniand a hydrocarbon or'mixture of hydrocarbons as above described at elevated temperatureoati leastSSOqF; andpreferably 900 to 1100? F. while maintaining conditions conducive to isomerizationreactionsas contrastedv with excessive dehydrogenation and crackinglreactions. Itis important to maintain in thereaction zone a substantial amount of hydrogen," at least 500 standard` cubic feet ofv hydrogen per barrel of charge stock.' The hydrogen 'separated from the product will usually be recycled tothe reaction zonerwith or withoutzmake-up hydrogen. The reaction pressure shouldbe suicientlyihigh to substantially eliminate paraiin dehydrogenation.- Thus, the reaction pressure should beatleast 50 pounds per square-inch and preferably about` lOGfto about 2000 pounds per'square inch. (These pressures-are gauge pressures as vare'other pressures mentioned in the specication andclaim.) The etects ofhydrogen concentration and reactionV pressure are related and'it is preferred `to use higher pressures when employing the lower hydrogen concentrations.

We have discovered'in accordance withv the invention' that byintroducing a minor-'amount of anoxygenic substance into the reaction Zone with the charge stock'and hydrogen, the nature of theY reactions taking placeis altered so that the extent loffisomerization is substantially increased while dehydrogenation and cracking of parans aredecreased asl compared with therelative extent ofV these reactions'takiug: place under thersame conditions but without the-addition'ofan' oxygenic substance. We have also discovered that in operations designed to produce a' selected yield of liquidproduct, the temperature in the present 'process should be maintained at least about '50 F. above the temperatureatwhich an equal amount ofliquid product'wouldV be obtained when subjecting the charge stock to identicalreaction conditions but omitting the use of the oxygenic substance.

The catalysts are the oxides of metals of groups VA,

ViA, and Vlil of the periodic table, such as molybdenum,

tungsten, cobalt and nickel. The most important of these-"metals is molybdenum and the description of thel invention hereinafter will be largely restricted to the useof oxides of molybdenum as the catalyst. While the catalyst metal oxides can b'e usedalone, they are preferably used'incombination wth carriers such as activated alumina, alumina gels, peptized alumina gels, silica gels, silica-alumina gels, aged or deactivated silica-'alumina cracking catalysts, silica-magnesia gels, magnesia, titania, bauxite,l and the like. A particularly suitable support comprises alumina stabilized with about 5` percent by weight silica.

The results of lour procedure of high temperature hy-v droisomerization While adding an oxygenic substance to the conversion-zone as compared-'with results of a similar operation in-Which no oxygenic substanceis added include markedly improved conversion to desired products and high desired product yields. The superior results of our procedure are obtained with all of the suitable charge stocks. Thus, the conversion of straight chain parains to the branched'chain isomers is considerably improved while total liquid yields' are substantially asgood or improved, and in the treatment of parainic naphthas our process in addition to high conversion of straight chain parafns to branched chain isomers also shows good conversion of naphthenes to aromatics and high liquid yields. It will be understood that in referring to liquid products, We mean butanesrand heavier.

While it is not intended to limit the invention to any theory, the research work in connection with vthe invention has indicated that the inclusion of a minor amountl of an oxygenic substance in the'total charge to the reaction zone has aneect on the statev of oxidation of the metal oxide catalyst so that the catalyst is maintained in possibly a lower oxide. Under conventional hydroformmg conditions, presumably the catalyst undergoes no furaatnte-d. Dec. t 1e, 195s,-

3 ther oxidation but rather additional reduction. By including a minor amount of an oxygenic substance in the charge to the reaction zone in accordance with our invention, it is believed that conditions are created in the reaction zone conducive to the conversion and/or maintenance of at least the surface layer of the molybdenum catalyst substantially in the form of M002. It is frequently desirable to include a minor amount of an oxygenie substance in the hydrogen employed for reducing the catalyst prior to the conversion reaction. When this is done, the prereduction will proceed only so far as to convert the surface of the catalyst to a state of oxidation corresponding substantially to M002. The use of the oxygenic substance in the charge to a reaction zone containing a catalyst prereduced in this manner will accomplish primarily maintenance of the desired gross state of oxidation of the catalyst. The results obtained in the practice of the invention show that when at least the surface of the catalyst is substantially in the form of M002, the catalyst is particularly effective as a hydroisomerization catalyst at elevated temperatures. Accordingly, the invention, when involving the use of a molybdenum catalyst, can be dened as carrying out a hydroisomerization process while maintaining at least the surface of the catalyst substantially in the form of M002. When catalysts in general are considered, it can be defined as comprising maintaining at least the surface of the catalyst in a gross state of oxidation intermediate the free catalyst metal and the highest oxide of the metal.

The term oxygenic substance is used herein in its usual sense and means oxygen and oxygen-containing compounds. It will be understood that an oxygenic substance can be employed which has no elect on the isomerization catalyst other than oxygenating it or substantially stopping its further reduction under hydroisomerization conditions. Among the many suitable oxygenic substances that can be mentioned are: oxygen; the inorganic oxygen compounds, steam, and the oxides of carbon or nitrogen; organic oxygen compounds such as the alcohols, carboxylic acids, aldehydes, esters, and ethers; and various oxygen-containing solids, for example, the catalyst containing the metal in a higher state of oxidation than is maintained in the reaction zone. Obviously, only such oxygenic substances should be used as have no detrimentalelfect on the catalyst or on the chemical reactions of hydroisomerization. Thus, it would ordinarily` be undesirable to use sulfur-containing oxygenio substances such as the oxides of sulfur, sulfuric acid, sulfonic acids, or the like, because of the poisoning effect of sulfur on the catalyst. substances, the most suitable are free oxygen, steam, carbon dioxide, andthe oxidized catalysts, because all of these substances are readily available and produce no complicating side reactions. However, if the reaction is carried out using a xed solid bed of catalyst, it would not be feasible for obvious reasons to employ the oxidized catalyst as the oxygenic substance.

We have indicated previously that the reaction temperature `for optimum results should be at least about.

50 F. higher in the present process than in a similar process `in which the state of oxidation of the catalyst is not controlled. As indicated above, the molybdenum in a hydroisomerization catalyst following oxidative regeneration which is ordinarily employed after a period of hydrocarbon conversion, exists substantially entirely as molybdenum trioxide. The catalyst in this state is too inactivefor optimum isomerization but when placed on stream and contacted with the charge of hydrocarbons and hydrogen, the catalyst is reduced to molybdenum dioxide and to a certain amount of lower oxide and/or free molybdenum having excessive cracking activity. Also, when the regenerated catalyst is reduced prior to contacting it with charge, it also reaches a stage of reduction characteristic of excessive cracking activity. Conversion at elevated temperatures. with the catalyst Of all the mentioned r in this state' results in low liquid product yields. We

have discovered that the excessive cracking activity whichthe molybdenum catalyst develops when contacted with hydrogen can be avoided by introducing a minor amount of an oxygenic substance to the reaction zone during the on-stream period. Thus, while by conventional tests the catalyst appears less active at the same operating conditions, the decrease in over-all activity is .largely due to the elimination of hydrocracking activity. Accordingly,.-

where an oxygenic substance is employed in the reaction and maximum isomerization is desired, the reaction is preferably carried out under reaction conditions including a substantially higher reaction temperature than employed in the absence of an added`oxygenie substance.

l'n many instances the temperature should be substantially higher in our process than 50 F. above the temperaturc required for the same liquid yield in the absence of an ougenic substance. Thus, in the hydroisomeri1 zation of normal pentane with the introduction of oxygen to the conversion zone in accordance with our procedure, the temperature is preferably at least about 200 F.` higher than the temperature required for best results in hydroisomerizing normal pentane without adding oxygen. This is surprising inasmuch as the hydroisomerization of normal pentane in the absence of added oxygenic substances at a temperature 200 F. above the optimum temperature for such hydroisomerization would result in substantially complete gasification of the charge. As stated above, the operable temperature range for our process in general is between about 850 and about ll00 F., with the temperatures at the lower end of the range being suitable for the heaviest stocks to which our process is applicable. O

The hydrogen concentration in the present process 1s maintained at about 500 to about 30,000 cubic feet of hydrogen per barrel (42 U. S. gallons) of liquid hydrocarbon charge and preferably at about 1,000 to about 20,000 cubic feet per barrel. Initially in the process it may be necessary to supply this hydrogen fromv an extraneous source, but since the reactions occurring during the reaction period are generally not hydrogen consuming or are hydrogen productive, as the reactions proceed, the hydrogen requirements can usually be met predominantly or entirely by recycle of hydrogen from the product.

The process of the present invention can be carried out according to several alternative methods of operation. For example, the process can be carried out with the catalyst in a stationary xed bed, in which case the catalyst is usually in the form of granules or pellets. When using the catalyst in a xed bed in an operation where periodic regeneration is practiced, the reaction cycle, including the on-stream period and the regeneration period, is carried out on the catalyst in the reaction zone.

The process can also be carried out utilizing a iinely divided catalyst in a uidized state. While in this case the on-stream and regeneration operations also can be both carried out with the catalyst in the reaction zone, it is generally preferred to provide a separate regenerator to which catalyst from the reactor can be conveyed continuously or periodically. The regeneration can be accomplished at about the pressure in the reactor or at lower pressures, for example, atmospheric pressure.

The space velocity can be between about 0.25 and 10.0 volumes of hydrocarbon per hour per volume of catalyst (based on packed catalyst volume). Generally a space velocity of about one-half to three vol./vol./hour is preferred.

Where the over-all effect of the reactions taking place is endothermic, provisions are made for supplying heat to the reaction zone during the on-stream period, such as by preheating the hydrocarbons, hydrogen, and catalyst to a suitable temperature.

There are a number of possible methods of introducing ethane, and` propane) fthe-required oxygenic `substance vtotheyreaction 'zone' in accordance with `our -invention. "Thus, using "pxygen gas asan-example, asmall quantity'ofoxygen can be mixed directly with the hydrogen'recycle'stream priorto its passage through the-preheaterzand before its mixture with the hydrocarbon charge. In this case precautions `should be taken to avoid explo'slonias'ffbywmixing a large proportion of an inert .gassuchfasadryt-gas.(methane, @with the oxygen'ibefore adding it to the hydrogen. Another possible procedure is to introduce the oxygen directlyvinto Vthe: catalytic reactor. In the case of a fluid catalytic reactor; the oxygen preferably is introduced at the bottomofthe-reactor so-as to facilitate its contacting the entire catalystbed. *In-:the case of a stationary lixed bed catalyst, it may 'b e necessary to introduce the oxygenI atf-a-number of points so as to achieve uniform treatment of the entire bed. .Still another method of introducing at 1east.a .part of the required oxygen to the catalytic reactor is todissolve oxygen in the liquid hydrocarbon charge.

A method of introducing oxygen to'the reactor which can be used in moving bed iluid catalyst processes is the incorporation of oxygen inthe catalyst transporting gas which is used to transport catalystfrom the regenerator to the reactor.

Also in moving bed fluid catalyst processes the oxygen can be supplied to the reactor by introducing a part of the oxidized catalyst from the Vregeneratorinto the reactor without an intervening or with only partial intervening reduction. The result is that the oxidized catalyst upon reduction by the hydrogen in the reaction zone produces the desired oxygenating conditions. 11n achieving this result a particularly effectivey procedure is to -divide the stream of regenerated oxidized catalystowing from the regenerator and to by-pass any reducing zone between the regenerator and the reactor withone of the streams so that a selected portion of the catalyst is reduced and a selected portion is not reduced, with the unreduced catalyst being charged to the reactor in the amount necessary to yield the proper amount of oxygen.

We have stated that the catalyst following regeneration can be placed on stream inthe hydroisomerization process either with or withouta prereduction. 'However, in the event that no reduction is employed, Vstarting the reaction with a completely oxidized catalyst will result in an initial period of uneven reaction conditions. There is also a short period of uneven conditions even if the catalyst is prereduced with hydrogen, during which the catalyst assumes its desired equilibriumstate of oxygenation. Also, the catalyst prereduced `with hydrogen may not reach optimum activity because the oxygenic substance may only retard the further reduction and mayV not accomplish reoxidation to the desired ;=statev ofoxidation. A high satisfactory procedure for eliminating either of these possibilities of an uneven initial reaction period is the reduction of the catalyst before its `being placed on stream with a ymixture of hydrogen and an oxygenicV substance preferably in the proportionsusedin the on-stream period. In this way the catalyst can be-adjusted to its desired equilibrium state of oxidation before being placed on stream, and uniform reaction conditions can be maintained throughout the cycle. VThis procedure of prereducing the catalyst lwith a mixtureof vhydrogenand .oxygenic substance is applicable to .either the vixed bed or moving bed procedures, whether using stationary or fluid type catalysts.

A particularly suitable method of maintainingvthe proper oxidizing conditionsv lin the reaction zone is to recycle the steam formed in the reactor sothat addition of fresh oxygenic substance is unnecessary after the initial equilibrium is achieved The recycle of steam in thegproper amount can be accomplished sbyzmaintaining the hydrogen to be recycled at a sufficiently hightemper ature to prevent condensation -of'the desired concentration ofwater vapor, Vand returning the Water vaponto the reaction zone with the .recycledhydrogen If a higher content of oxygenic substance is desired, fresh oxygenic vsubstance can be addedto the hydrogen recycle stream.

The beneficial efIects of the addition of -an oxygenic substance to the hydroisomerization reactor on a molybdenum catalyst are obtainable at a very low concentration of oxygen in the reaction gases. Inl general 'the maximum beneficial effect is obtained when the oxygen content is about 0.1 to 1 volume percent of the hydrogen. Some beneiit is obtained,however, when the oxygen content is of the order of 0.05 Volume percent of the hydrothe extreme lower limitgand the preferred lower limit,

respectively. vThis means that'an oxygenic gas such as steamawhich per mol contains only 0.5 mol of molecular oxygen must be used in double the volume required for pure oxgen. When the hydrogen contains about 0.075 mol of oygen per mol ofhydmgen, hydroisomerization is decreased although high liquid yields are still obtained. Larger amounts of Voxygen further decrease the value of the addition. The preferred upper limit corresponds to about 0.03 mol of oxygen per mol of hydrogen. Actually in the case of oxygen, it is generally not advisable to add substantially more than about 0.01 mol of oxygen per mol of hydrogen, because in such case the safe composition for a non-explosive mixture of hydrogen and oxygen may be exceeded. The higher concentrations of oxygen in hydrogen are only obtainable by use of oxygen equivalents such as steam. If steam is usedas the oxygenic substance, the maximum concentration of the oxygenic substance may be affected by the steam stability of the catalyst which is being used. Generally, a steam partial pressure of about 50 pounds per square inch should not be exceeded so that when the reaction pressure is high the steam stability of the catalyst could govern the maximum concentration of oxygenic substance which could be used.

Sulfur has a detrimental eiect on the catalysts in our process and although there is some variation in sulfur tolerance, it can be said in general that the beneficial effects of the addition of oxygen are lost or greatly decreased when ,the catalyst is substantially sulded. The molybdenum catalyst, for example, drops considerably in activity .when the molybdenum is about 50 percent sulfided, i. e., when the ,catalyst contains about 50 percent of the amount of sulfur required to convert the molybdenum to molybdenum disulfide, and the benefit of operating with an oxygenic substance inthe hydrogen is lost. Consequently,

.it is desirable to employ charge stocks of low sulfur contentin order to operate yfor long on-stream periods without regeneration. When treating a high-sulfur stock such as West Texas naphthaby our process, a prior desulfurization treatment may be required if long throughputs are desired, yalthough it maybe more convenient to have no prior desulfurization ,treatment and instead to regenerate the sulfided catalyst at frequent intervals. When the process uses a catalyst which is damaged by regeneration, the suliding should -be avoided by the prior desulfurization of charge stocks having high sulfur contents.

We have lconducted hydroisomerization runs in accordance with the process of our invention which show thev superiority 1of our process,..whether .petroleum fractions or` otherfrhydrocarbonsvare used-as charge stoeks, over-processesin whichnooxy-genic substance `is added 7 to the conversion zone. l The details of procedure were as follows:

EXAMPLE 1 A molybdena-on-alumina fluid catalyst consisting of 10.8 percent by weight molybdenum trioXide deposited on activated alumina coprecipitated with about weight percent silica was calcined in air at 1100 F. for several hours and then reduced in a stream of pure hydrogen at 1050 F. for 4 hours. Then to hydroisomerize a 375 F.

Cit

VEXAMPLE 3 The procedure of Examples 1 and 2 was repeated using a reaction temperature of about 890 F. The hydrogen used in the prereducton and the on-strearn period conend point West Texas straight run naphtha having the inl() t@inet-l 025 Volume Demont oxygen The results are .spection data listed in Table I below, the reduced catalyst recorded 111 Table I beloW- was brought to reaction conditions including a tempera- We have also Collductefl 1115 111 the absence of added ture of about 968 F., a pressure of about 300 pounds per oxygen for Comparison Wltll the results 0f 011r Processvsquare inch gauge, a space velocity of about 1.0.volume 1 The detalls of Procedure Were aS follows: of liquid naphtha per hour per volume of catalyst (based EXAMPLE 4 on densely packed catalyst volume), and a hydrogen concentration of about 20,000 standard cubic feet of hydrogen h per barrel of naphtha over a throughput of about 1.0 treri gaat gli lgg svilrg lslnlacrtlle 5% 1/a; volume of naphtha per volume of Catalyst' Following 70 tem erature of about 875 F but usin subsilantifll this lining-out period, product was collected for an on- OX pen free h dro en (i e cntainnn 155s than abmi; stream throughput of about 2.0 volumes of naphtha per o Oylg volume yercegnt ox'ne'rl) for the rereducton and volume of catalyst The hydrogen used in the Hning'out Vthe on-streamp eriod li ou enic sulstances were in and Orl-Stream periods Contained Oxygen in the amount troduced to thepconversion zonxegdurin the rocess The of about 0.44 percent by volume. The results of this run 2r results are recorded in Tabl I below g p in terms of product yield and characteristics are listed in d e Table I below. EXAMPLE 5 EXAMPLE 2 The procedure of Example 4 was repeated using a tem The molybdena-alumina fluid catalyst of Example 1 l perature of 840 F. The results are recorded in Table I was calcined and nitrogen purged and then reduced in a o below.

Tablel Example No 1 2 3 4 5 Temperature, F 968 924 800 875 840 Pressure: p. s. i. g.. 300 300 300 300 300 On-strearn throught Vol./Vo1- 2 2 2 2 2 Space velocity, Vol./Hr./Vo1 1 1 1 1 1 Space velocity, WtJHn/W't. 0.76 0. 76 0.76 70 0. 7G Hydrogen rate, STP Cu.Ft./Bbl 20, 300 19, 200 20, 100 20, 000 10, 700 Oxygen in Hydrogen stream, Percent by V01 0. 44 0. 0.25 0. 01 0. 01

Recovery: Percent by Wt. of Liquid Charge:

Liquid product 71. 5 78. 9 84. 1 '15. 0 83. t Dry gas 14. 9 s. 7 G. 6 13.1 o. o Wet gas 7. 5 6. 7 4. 8 5. 5 3. 7 Carbon 0.3 0. 3 0.3 0. 3 0. 3

Total 94. 2 94. e 95. s 93. 9 99. 4

Recovery, Percent by Vol. of Liquid Charge:

Liquid product 69. 5 77. 7 s3. 4 75.0 s3. a 10 RVF product 74. 7 om 9s. 5 s1. 3 93. t

Charge 4s. 5 50. 2 19. 9 0. 7561 o. 778s o. 7500 o. 009 o. 014 0.010 1.0 1.6 0.8 Bromine No 1. 9 l. 3 2. l 1. 0 `.Aromatic Content, Percent by 13. 9 63. 1 52.8 45.5 47. 0 41. 7 Vapor Pressure, Reid, Lb 0 6. 0 5. 4 3.7 5. 7 3. 7 Distillation, Gasoline- Over point, F 270 109 104 120 114 130 End Point, F 315 417 409 s 38e 401 10% at, "F 297 149 15s 139 165 197 at, F 316 26s 282 296 261 289 o a 0F 343 336 340 344 32s 3:10 Octane No., Ilcro Research Method- Clear 34. 8 97. 0 91. 5 83. 3 87. 2 78, +3 ce. TEL 52.8 106.1 98. 4 94. 5 95. s 91. Properties of 10 RVP product:

Gravny, API 49. 2 51. s 54. 5 54. 4 55. Olens, Percent by 0.8 1.1 0. 0 1. 5 0. Aromatic Content, Percent by Vol 5g, 7 4g. 6 40 6 44. 2 37. Octane No., Micro research method:

97. 2 92. 2 85. 1 ss. 2 s1. +3 ce. TEL l106.3 99.1 96.3 97.8 91. Butaues:

Percent by Vol. required for 10 RVP 17. 4 6. 9 7.9 10. 8 7. 9 10 Percent by Vol. available 2 0 12. 9 10. 2 6.7 B. S 5,

1 Performance number.

2 Based on volume of 10 RVP gasoline.

stream of hydrogen Acontaining 0.34 volume percent oxygen for 4 hours at 1050 F. The catalyst was then Table I clearly shows the superior results obtainable by the process of our invention in terms of gasoline yields placed on-stream to hydroisomerize the West Texas and octane ratings. An even clearer representation of the improvement is aflordedfby Figure 1", which plots the data froml Table I for liquid product yield and research octane numbers, clear, of the total liquid product, adjusted to pounds Reid vapor pressure by the addition of butanes.

Referring to Figure l of the drawing, curve A plots the data for Examples 1, 2, and 3, and curve B plots the data for Examples 4 and 5. It is clear from a study of these curves that for any particular liquid product yield the octane rating of the product would be higher for the process of our invention (curve A), and likewise for any particular octane number the yield would be higher. Thus, the curves indicate that for a clear octane rating of 88.2 the yield for the process of our invention would be about 90.0, as compared with only 81.3 percent for the process carried out without the addition of oxygen. Comparable superiority in yield for our process would be shown for any other desired octane level.

We have conducted additional runs involving hydroisomerization of a charge consisting of a single hydrocarbon in which oxygen was the oxygenic substance ernployed. Details of the procedure were as follows.

EXAMPLE 6 A iluid catalyst consisting of 8.4 percent molybdenum trioxide deposited on activated alumina impregnated with about 4 percent silica was calcined and then reduced in a stream of hydrogen containing about 0.3 mol percent oxygen at 1050" F. for 4 hours. The reduced catalyst was placed on-stream for hydroisomerization of normal pentane under reaction conditions including a temperature of about 875 F., a pressure of 300 pounds per square inch gauge, a liquid hourly space velocity of about 1.0 (based on packed catalyst volume), and a hydrogen concentration of about 21,000 standard cubic feet of hydrogen per barrel of liquid pentane. During the reaction the hydrogen stream contained about 0.3 mol percent oxygen. The product of the throughput interval from 1.0 to 3.0 volumes of charge per volume of catalyst, inclusive, was collected. The yield of liquid product for this interval was 86.1 percent by weight and the product contained 3.2 percent by weight isopentane.

EXAMPLE 7 The procedure of Example 6 was repeated at a reaction temperature of about 925 F. The yield of liquid product for the throughput interval from 1.0 to 3.0, inclusive, was 78.6 percent by Weight, and this product contained 11.1 percent by weight isopentane.

EXAMPLE 8 EXAMPLE 9 The treatment of normal pentane according to the procedure of Example 6 was repeated but using a hydrogen stream for both the prereduction andv on-strealn periods which, was substantially free of oxygenc. substances, and employing a reaction temperature of about 721 F. The yield of liquid product of the throughput interval from 1.0 to 3.0, inclusive, was only 51.5 percenty by weight, and

this product contained only 0.5 percent by weight isopentane.

EXAMPLE 10 The treatment of normal pentane in accordance with Example 9 was repeated at a reaction temperature of about 772 F. The yield of liquid product for the throughput interval from 1.0 to 3.0, inclusive, was only 41.1 percent by weight and contained only 1.8 percent by weight isopentane.

The results of Examples 6 through l0 illustrate very clearly the superior results of. our process, since for all runs conducted in accordance with our process (Examples 6, 7 and 8) both the liquid yields and the percentages of isopentane in the product were much higher than for the process in which no oxygen was introduced to the reaction zone. These examples show also that a considerably higher temperature is employed in our process. Thus, Example 10, which was carried out at a temperature of 772 F., gave a very low liquid yield and indicated that a higher temperature without the addition of oxygen would be undesirable. The lowest temperature run in accordance with our process (Example 6) was about F. higher and was not necessarily the highest optimum temperature, since the isopentane yield was not as high as was obtained in the higher temperature runs.

We have also conducted runs which compare the results of hydroisomerizing methylcyclopentane `by our. process using oxygen as the oxygenic substance, and by the process in which no oxygen. is added. Details of the procedure were as follows.

EXAMPLE 11' EXAMPLE 12.

The hydroisomerization of methylcyclopentane accord'- ing to the procedure of Example 11 was repeated using a reaction temperature of about 1002 F. The yieldy of liquid product for the throughput interval from 1.0 to 3.0, inclusive, was 88.2 percent by weight, and this product contained 20.6 percent by volume. benzene.

EXAMPLE 13V Methylcyclopentane wasV treated at a temperature of about 848 F. according to the procedure of Example 11 but usingz a hydrogenstream for. both the prereduction and on-stream periods which was.. substantially free. of oxygenic substances. The yield of liquid product for the throughput interval from 1.0 tol 3.0, inclusive, was 86.5 percent by Weight, and this product contained 1.9 percent by volume benzene.

EXAMPLE 14 T heY treatment of methylcyclopentane according to the procedure of Example 13 was carried out at about 896V F. without the addition of oxygen. The yield of liquid" product for the throughput interval from 1.0'to 3.0, in-

elusive, was 65.6 percent by weight, and this product contained 4.2 percent by volume benzene.

The results of the methylcyclopentane isomerizations can be evaluated in terms of benzene production, the benzene being produced by isomerization of methylcyclopentane to cyclohexane which dehydrogenates to` benzene. The results clearly show the superiority of our process as carried out in Examples 1l and 12. Examples 11 tor14 show that benzene content'of the product increases with temperature while total yield of liquid product decreases with temperature. Examples 12 and 13 produced comparable yields of liquid product although the yield for our process (Example 12) was somewhat higher. The signicant fact, however, is that for comparable liquid yields, our process produced more than ten times the amount of benzene produced in the process carried out without the addition of oxygen (Example 13). The results of Example 11 show that when the temperature is somewhat lower than in Example 12, the total production of benzene is somewhat lower although still much higher than produced by either Examples 13 or 14, and the liquid product yield is higher. Example 11i shows that when the temperature is increased above that of Example 13, the yield of benzene is somewhat better, but the total liquid yield is considerably lower, so that obviously no temperature could be selected for the process carried out in the absence of oxygen which would show both the high liquid yields and high conversion 1o benzene of our process.

We have also conducted experiments which demon strate the applicability of our process for hydroisomerizing a synthetic mixture of a normal paratlin, a naphthene, and an aromatic compound. The procedure was as follows:

EXAMPLE 15 The molybdena-on-alumina uid catalyst consisting of 8.4 percent by weight M003 supported on alumina impregnated with about 4 percent silica was precalcined at 1100 F. in a stream of air and then prereduced at 1050 F. in a stream of hydrogen containing 0.28 mol percent oxygen. The catalyst was then flushed with pure nitrogen, cooled to about 900 F., and pressured to about 300 pounds per square inch gauge with hydrogen containing 0.28 mol percent oxygen. This catalyst was then placed on-stream for hydroisomerizing a mixture consisting of 39.1 percent by weight pure-grade normal pentane, 48.7 percent by weight pure-grade cyclohexane, and 12.2 percent by weight chemically pure benzene. Reaction conditions included temperature of 899 F., space velocity of 1.0 volume of liquid charge per volume of packed catalyst per hour, pressure of 300 pounds per square inch gauge, and hydrogen concentration of 21,000 standard cubic feet of hydrogen per barrel of liquid charge. The hydrogen stream contained oxygen in the amount of 0.28 mol percent. The product of the throughput interval from 1.0 to 3.0, inclusive, was analyzed and the results in terms of the yield and characteristics of this product are given in Table II below.

EXAMPLES 16-23 Mol Percent Oz in Ha Stream Temperature, Example No. F.

In each of the above runs the catalyst was prereduced with a vhydrogen stream containing the same amount of EXAMPLE 24 The hydroisomerization of the normal pentane, cyclohexane, benzene mixture of Example 15 was repeated using the catalyst of Example 15, precalcining and prereducing the catalyst in the manner described with the exception that the hydrogen stream for prereduction contained steam instead of oxygen in the amount of 17.1 mol percent steam. Reaction conditions included temperature 960 F., space velocity 1.0 volume of liquid charge per volume of packed catalyst per hour, pressure 300 pounds per square inch, hydrogen concentration 21,000 standard cubic feet of hydrogen per barrel of liquid charge, and the hydrogen recycle stream contained 0.8 mol percent steam. The results in terms of the yield and characteristics of the product of throughput interval from 1.0 to 3.0 inclusive volumes of liquid per volume of packed catalyst are listed in Table Il below.

EXAMPLES 25-26 The procedure of Example 24 was repeated, but using hydrogen for prereduction which had the same steam content as the on-stream hydrogen, i. e., 0.8 mol percent, and using temperatures of 975 F. and 985 F., respectively. The yields and characteristics 0f the products for the throughput intervals from 1.0 to 3.0 inclusive are given in Table II below.

The superiority of our procedure for hydroisomerizng the n-pentane, cyclohexane, benzene mixture is shown by comparative runs on this mixture without the addition of an oxygenic substance to the reaction Zone. Details of procedure were as follows:

EXAMPLE 27 The molybdena-alumina catalyst of Example 15 was precalcined and then prereduced in a stream of pure hydrogen and placed on-stream for the treatment of the n-pentane, cyclohexane, benzene mixture. The hydrogen stream was substantially free of oxygenic substances and no oxygenic substances were added to the reaction zone after the start of the process. Reaction conditions included temperature 851 F., space velocity 1.0 volume of liquid charge per volume of packed catalyst per hour, pressure 300 pounds per square inch gauge, hydrogen concentration 21,000 standard cubic feet of hydrogen per barrel of liquid charge. The yield and characteristics of the product for the throughput interval from 1.0 to 3.0 volumes of liquid charge per volume of packed catalyst inclusive are listed in Table II below.

EXAMPLES 28-32 The procedure of Example 27 was repeated with variations in the reaction temperature as follows:

The yields and characteristics of the products for throughput intervals from 1.0 to 3.0 volumes of liquid charge per volume of packed catalyst inclusive are listed in Table II.

Of Ca, percent i by, Wt.

-reducoxida- 7 .995730254749822 4 m44.7.6.2.6.5.5.2. .8. 3 3338333333M2WM gure 3 shows a percent Cyclohexane lie. 5

procedure of'hydrog a ,normal-par- As'Table II and ployed as ,theoxygenic concentration,A of steam Methylcyclopentime total pentanes with increase in process, while curve B shows a Composition Y of I :cyclohexane and saturatodC isomers, 100 Vol.- rpercent EXAMPLES 3 3-3 7 Cyclohexane and sat- Hexane urated Ce `isomers s, percent Benzene

pentanes in our process and in'the process in which no oxygen is used. Thus, curve` A of Fi sustained high yield of isopentane yield for our arp drop in total pentane yield as ythe yieldof iso- It is-quite clear from the results shown in Tablel II andn Figures 2 and 3 that our tion conditionsin the reaction zone is a great improve- We have conducted additionalruns whichV `show the by Wt.

sh pentane increases.

isomerizing underV properly lcontrolled ,oxidation ment over operating without proper control -of the tion-reduction conditions when treatin ahn, naphthene, aromatic mixture.

Figures 2 and 3 show, our superior results were obtained witheither oxygen or steam em substance although a higher mol was used.

n n o effect on hydroisomerlzanon results of'the concentration of the oxygenic substance in the conversion zone. tails of procedure were as follows:

Table'II rge ere

`Yields of Various Product ygenic Liquid Product,

t Charge Mixture vEven in these two g. Thus, re-

Run

ure of hydroisomerizing um catalyst in the proper Thesuperior results of Mol percent Oxygenic substance in Hydrogen Stream it is seen that our process, as repreloives sustained high liquid product Charge Example No.

ments afforded by our proced while maintaining the molybden state ofoxidation by theA introduction of an oxygenic substance to the reaction zone.

our procedure. as compared with treating the same cha stock without the addition of an oxygenic substance W obtained with either steam or oxygen as the ox substance.

oxygen was added (Example 27).

latter runs the total liquid product yields were substantially equal while the yield of isopentane was much greater for our process and the yield of benzene was somewhat greater.

sented by curve A,

The n-pentane, cyclohexane, benzene mixture described in Example 15 was subjected to hydroisomerization under conditions including temperature of about F., pressure of about 300 pounds per square inch,

pace velocity of 1.0 volume of liquid charge per Volume of packed catalyst per hour, and hydrogen concentration of 21,000 standard cubic feet of hydrogen per barrel Runs 33, 34, and 35 were made using -alumina uid catalyst consisting of 8.4

percent M003 on alumina impregnated with about 4.

O C OO f yields up to a temperature above 1000 F. while the liquid yield for the process in which no oxygenic substance was used drops sharply at a temperature of about 870 F. as shown by curve B. Thus, if temperatures suciently high to give the desired high conversions to s branched chain parains are employed without the addition of oxygenic substances, cracking is excessive and liquid yields decrease greatly.

Figure 3 gives a clear comparison of the relationship between yield of isopentane and recovery of total Total yield of liquid products, percent by Wt.

Tre

percent by wt.

Total yield 0f C5, percent by Wt.

Cyclohexane Benzene and saturated Cs isomers Table III Il-Cs I1-C5 Yield of Various Products, percent by Wt.

C1 t0 C4 Run equivalent inhydrogen Temp., u

M o1 percent steam equivalent used for pre1-eduction Mol percent steam m onstream hydrogen 1 Actually 0.15 O2.

Example No.

Runs 36 and 37 were made with they:Y

cent M003 on alumina coprecipitated with to 6 per-A f cent silica. The concentration of oxygenic substance in .A

the hydrogen used for prereduction and the hydrogen used in the on-stream period was different for each run.

The concentrations used are listed in Table III. Runs 33, 35, 36, and 37 used steam as the oxygenic substance and run 34 used oxygen. The table lists the steam equivalent of the oxygen concentration used in Example 34. Table III also lists the yields of products obtained during thethroughput interval of from 1.0 to 3.0 volumes of liquid charge per volume of packedl catalyst inclusive.

The data of Table III show that the yields of benzene and isopentane for Example 36 were considerably better than the yields of the same products for Example 37 so that a steam concentration of 1.0 mol percent (Example 36) appears to be better than aesteam concentration of 13.2 mol percent (Example 37). The concentration of steam in Example 37 appears to have been so high as to reduce the catalysts isomerization activity although it should be noted that a very high liquid yield was obtainedj in Example 37. It is also noted that in Example 33, the

total liquid product yield was lower than in any of the other runs and there was a high yield of gas (C1-(2.,) so that apparently the cracking activity of the catalyst was not thoroughly suppressed by 0.14 mol percent of steam in the hydrogen stream. These data support our previous remarks concerning the proper concentration of oxygenic substances.

As stated previously, the present process is particularly effective when employed for the treatmentof hydrocarbons boiling in the naphtha and gasoline boiling range; i. e. hydrocarbons which in admixtureproduce a mixture boiling within the range of about 32 to about 550 F. Thus, the hydrocarbons can be referred to as naphtha hydrocarbons.

The detailed description given above has been largely concerned with operations involving the use of molybdenum catalyst, but in general the conditions disclosed as being suitable for these operations will also be suitable when another catalyst of the class disclosed above is used. However, in many cases the optimum amount of oxygenic substance to be'added will vary somewhat depending upon -the particular catalyst employed.

We have found for example that the concentration of oxygenic substance in the hydrogen to suppress cracking when a catalyst, comprising nickel oxide impregnated on alumina containing about 4 percent by weight of silica, is employed is higher than if a molybdenum ovide catalyst is employed and is a function of the amount of nickel in the catalyst. A steam concentration of about 1 mol percent in the hydrogen is necessary to substantially suppress cracking when the catalyst contains about 2 percent by weight nickel oxide. Higher steam concentrations are necessary when the catalyst contains higher concentrations such as 5 percent by weight.

Obviously many modifications and variations of the invention as hereinabove set forth may be made without departing from the spirit and scope thereof, and therefore only such limitations should be imposed as are indicated in the appended claim.

We claim:

A hydroisomcrization process which comprises reducing a regenerated molybdenum oxide catalyst with hydrogen containing an amount of an oxygenic substance selected from -the group consisting of oxygen, steam and carbon dioxide supplying about 0.0023 to about 0.03 mol of oxygen per mol of hydrogen, contacting the resulting reduced `catalyst in a reaction zone with a charge comprising naphtha and hydrogen in proportions of about 1,000 to about 20,000 cubic feet of hydrogen per barrel of said naphtha and an amount of an oxygenic substance selected from theV group consisting of oxygen, steam and carbon dioxide supplying about 0.0023 to about 0.03 mol of oxygen per mol of hydrogen, at a pressure of about to about 2,000 pounds per square inch and at a temperature of about 850 to about 1100 F.

References Cited in the file of this patent UNITED STATES PATENTS 2,399,927 Howes et al. May 7, 1946 2,424,636 Smith et al. July 29, 1947 2,433,603 Danner et al Dec. 30, 1947 2,487,563 Layng Nov. 8, 1949 2,598,642 Hurley May 27, 1952 2,661,383 Beckberger et al. Dec. l, 1954 2,718,535 McKinley et al Sept. 20, 1955