US3175016A

US3175016A - Selective oxidative dealkylation

Info

Publication number: US3175016A
Application number: US108688A
Authority: US
Inventors: Charles J Norton; Thurle E Moss
Original assignee: Marathon Oil Co
Current assignee: Marathon Oil Co
Priority date: 1961-03-20
Filing date: 1961-03-20
Publication date: 1965-03-23
Anticipated expiration: 1982-03-23

Description

March 1965 c. J. NORTON ETAL 3,175,016

SELECTIVE OXIDATIVE DEALKYLATION Filed March 20, 1961 6 Sheets-Sheet l PRESSURE CONDENSER PRESSURE 0- RELIEF VALVE WATER CONDENSER REF LUX) 5g coagggr ms AFFEL I m i; REACTOR P AIR FEED 3 5 1 g: gl g N N E! 5% 3 7 WET TEST s! 6 3 'ELT-C-I 5:. \7 DRY ICE TRAPS m: I 3i Fla-I FLUID-BED REAcmR INV ENT OR.

CHARLES J. NORTON BY THURLE E.MOSS

ATTORNE YS March 1965 c..:. NORTON ETAL SELECTIVE OXIDATIVE DEALKYLATION 6 Sheets-Sheet 2 Filed March 20, 1961 gma an :4 32 89 31 es- 3 0 3:5 32:30 282 3 :4 M a WIWH. 22 g a 3% m 8030A- aa-um 38 Ann Sa 3 2 2% 2.28 9:95 55.30

SEES-om C IN VEN TORS CHA RL ES J. NORTON THURLE E. MOSS ATTORNEYS March 1965 c. J. NORTON ETAL 3, 7

SELECTIVE OXIDATIVE DEALKYLATION Filed March 20, 1961 e Sheets-Sheet :5

Theoretical Limit 100 PERCENT NAPHTHALENE IN PRODUCT I I I I l I I I I I I I I I I I I I I I I I I I I I 16 PERCENT COMBUSTION 0F FEED FIG. 3 CATALYSTS FOR OXIDATIVE DEALKYLATIQN INVENTORS CHARLES J. NORTON THURLE E. MOSS AT TORNEYS March 1965 c. J. NORTON ETAL ,9

SELECTIVE OXIDATIVE DEALKYLATION Filed March 20, 1961 6 Sheets-Sheet 4 m g 90 Over 3.4 Wt. Gadmia-silica Catalyst g 80 e-i K E 10|ln 29 02 5'3 0 REACTION TEMPEATURE,

IN V EN TOR-S CHARLES J. NORTON THURLE E. MOSS BY ATTORNEYS March 1955 c. J. NORTON ETAL 5,

SELECTIVE OXIDATIVE DEALKYLATION Filed March 20, 1961 6 Sheets-Sheet 5 0 2-) a \d l 8553- ZE$OURS fr, l.0 LOG [wT CdO] 'HASE Oxmmow IF 1- IVER CADMIA AT 400C ELECTHVE VAPQR- Ki} 2 it} 21 E E Q h Li INVENTORS CHARLES J. NORTON THURLE E. MOSS ATTORNEYS March 1965 c. J. NORTON ETAL 3,175,916

SELECTIVEOXIDATIVE DEALKYLATION Filed March 20, 1961 e Sheets-Sheet e ISOYIELD CONTOURS FIG, 6 SELECTIVE VAFm; l T0 APMHALENE INVENTORS. CHARLES J. NORTON BY THURLE E.MOSS

ATTO RN E YS United States Patent 3,175,016 SELECTIVE OXIDATIVE DEALKYLATION Charles J. Norton and Thurle E. Moss, Denver, Colo, assignors to The Marathon Oil Company, Findlay, Ohio, a corporation of Ohio Filed Mar. 20, 1961, Ser. No. 108,688 Claims. (Cl. 260-672) This invention relates to the controlled selective vapor phase oxidation of alkyl substituted hydrocarbons to lower molecular weight hydrocarbons, and more particularly is directed to the control selective vapor phase oxidation for the dealkylation of alkyl substituted aromatic hydrocarbons to lower weight homologs of these compositions and specifically to produce the more desirable parent homolog of such compositions.

Any mixed hydrocarbon material, whether derived from petroleum, coal tar, shale oil, and similar materials, invariably includes in such a mixture a substantial portion of higher molecular weight homologs and derivatives of various of the hydrocarbon compositions, including such compositions as paraffins, olefins, aromatics and the like. The conventional recovery of aromatic hydrocarbons from such materials, whether the stream or mixture is a virgin or a processed stream or mixture, invariably includes a substantial amount of alkylaromatic homologs of the parent aromatic hydrocarbon, as well as other derivatives such as alkylhydroaromatic hydrocarbons of the monocyclic, bicyclic, or polycyclic homologous series. With treated streams, for example, treatments such as thermal cracking, catalytic cracking, hydro-forming, hydrocracking, etc., substantial amounts of alkylaromatic and alkylhydroaromatic hydrocarbons are found in the resultant mixtures.

The predominant homologous aromatic hydrocarbon series in a product stream depends upon the nature of the original stream, the treatment process and the severity of the treatment of the original stream. Furthermore, in most cases these streams are frequently richer in the higher molecular weight homologs or derivatives than in the parent homolog itself. Normally, such streams include a large number of isomers, at relatively low concentration each, which greatly complicates the separation, purification, utilization and marketing of individual members of the stream. For such reasons alkylaromatic hydrocarbons have not been as valuable commercially as the parent homologs, for example, benzene naphthalene, anthracene, etc., have been commercially more valuable as chemicals than their alkyl homologs.

The oxidation of hydrocarbons to oxygenated reaction products is well known in literature. Examples of such include the oxidizing olefins to epoxides, glycols, aldehydes, acids, etc., over various catalysts; alkanes have been oxidized by air or oxygen in the vapor phase, in the absence of catalysts, to mixtures of valuable oxygenated derivatives; naphthalene, methylnaphthalenes, phenanthrene, anthracene, tetralins and xylenes have been catalytically oxidized in the vapor phase over catalysts such as vanadium, molybdenum oxide, tungstic oxide, and various other mixed metallic oxides to phthalic anhydride, maleic anhydride, and/ or naphthoquinone, depending upon the composition of the feed mixture. Likewise, it has been known to oxidize naphthalene and ortho-xylene to phthalic anhydride over fixed or fluid bed vanadia; benzene and toluene have been oxidized to maleic anhydride, and toluene oxidized to benzyl alcohol, benzalde hyde and benzoic acid under slightly diiTerent conditions.

In the known oxidation processes of aromatic hydrocarbons, carbon dioxide, water, maleic anhydride, phthalic anhydride, naphthoquinone, and naphthol have been iso- 3, l 1 h Patented Mar. 23, 1965 lated as reaction intermediates or products from methylnaphthalene and naphthalene feed materials.

According to the present invention we have prepared naphthalene as a product from the oxidation of methylnaphthalene and have further oxidized other organic materials such as peroxide, ethers, alcohols, aldehydes, acids and esters, in the vapor phase under highly selective conditions to the parent hydrocarbon. Additionally, we have discovered a highly eificient two-step oxidation of methylnaphthalene to phthalic anhydride. By controlling the reaction conditions in relation to the particular catalyst, the oxidation according to the invention may produce naphthalene as a major reaction product from alkylnaphthalenes.

Included among the objects and advantages of the present invention is a selective oxidation process for the dealkylation of alkylaromatic hydrocarbons by controlled selective vapor phase catalytic oxidation over fluidized oxidation catalysts. The process provides a novel, efficient and economical method for the oxidative dealkylation of higher homologs in an alkylaromatic hydrocarbon series to lower molecular weight compositions and to the parent homolog of the series. The invention, furthermore provides a novel catalyst system for oxidative catalysis of various hydrocarbon systems.

Further objects and advantages of our invention relate to a novel process for producing parent homologs of aromatic hydrocarbons as well as a two-step vapor phase oxidation of certain aromatic hydrocarbon mixtures for the improved production of phthalic anhydride, maleic anhydride, and naphthoquinone. The process of the invention includes an efficient and effective process for producing naphthalene from various petroleum streams containing varying amounts of alkylnaphthalenes, and certain hydrocarbon streams enriched in alkylnaphthalenes derived from various refinery streams, whether virgin or treated steams.

Additional objects and advantages reside in the process for producing naphthalene from light catalytic cycle oil cuts which may be oxidized per se or from extracted portions which are selectively enriched in alkylaromatics.

Still other objects and advantages of the invention reside in modifying the conditions of oxidative cracking of hydrocarbon streams by the addition of controlled amounts of gaseous oxygen such as air and oxygen. Sulfur oxides and nitrogen oxides may be added to exert beneficial cracking effects far out of proportion to their stoichiometry and at reaction conditions of considerably less severity.

Additional objects and advantages will be readily apparent from the following description.

In producing the parent homolog of aromatic series, the feed material used for the oxidative cracking determines the ultimate products. Various types of feed materials may be used in the process and compositions such as alkylheterocyclic aromatic hydrocarbons may be oxidatively dealkylated to lower molecular weight homologs, similarly, organic derivatives such as alcohols, aldehydes, acids, ethers, esters, peroxides, etc. may be oxidatively reduced to the parent compound. In addition, crude mixtures containing the alkylaromatic hydrocarbons may be oxidatively reduced to produce the parent compound, and in one preferred process the aromatic hydrocarbons in a light catalytic cycle oil are selectively extracted from the oil to produce a hydrocarbon feed for the process.

Numerous materials may be employed as effective catalyst for the oxidative cracking. The principal metal oxides for catalysts may be one or more of the oxidation states of one or more of the metals of the periodic chart groups 1B, 113, I113, IVB, VB, VIB, VIII, IIIA, IVA,

VA and VIA, are preferably of the periodic groups IB, IIB, IIIA, IVA and VA, for example, silver, zinc, cadmium, mercury, aluminum, gallium, indium, thallium, silicon, germanium, tin, lead, bismuth, etc. Especially good selective activities were found for the oxides of silver, zinc, cadmium, indium, and bismuth. Surprising- 1y, cadmium compositions, and especially amorphous cadmium oxide, were found to be highly selective in the oxidative cracking process. In addition, compounds of cadmium and other metals of the group have been found to be highly selective, for example, cadmium silicate was highly selective in oxidative cracking.

The catalysts specified above may be modified by activators which may be chosen from the less selective but more active metal compounds of groups enumerated above, for example, chromium, manganese, iron, cobalt, nickel, and copper oxides. In addition, inihibitors under some conditions may be used to control the activity of the catalyst, and these may be chosen from various of the metal oxides or compounds of metals and non-metals in groups IA, IIA, IIIB, IIIA, and stable oxides or anions derived from the metals and non-metal groups IIIA, IVA, VA, VIA and VIIA under the reaction conditions. Such inhibitors or modifiers may also be alkali or alkaline earth metal salts of stable aluminates, silicates, phosphates, sulfates, for example, sodium, lithium or potassium sulfate, acid sulfate, or phosphate.

Such catalysts may be used alone or on a catalyst base, and it is preferably used as a fluidized catalyst bed. Such fluidized catalyst beds are well known in the art, and in general such beds include providing the catalyst material in small sizes and maintaining the catalyst in a fluid state by means of passing the vaporized hydrocarbon feed and air through the bed at a sufficient velocity to maintain the catalyst material suspended in the air.

The catalysts for use according to the invention should be dried and preactivated in air or oxygen at a temperature of about 500-600 F. for a suitable time. Spent catalyst may be regenerated under similar conditions. To maintain a continuous catalytic oxidative cracking reaction using recycled catalyst, the regeneration rate of the spent catalyst must be faster than the oxidative cracking rate so that the catalyst may be withdrawn, regenerated, and recycled into the reaction. This maintains the catalyst in a highly reactive state.

Supports for the various catalysts may be conventional catalyst supports, for example, alumina, silica, silicaalumina, silicon carbide, boria, beryllia, glass, ceramic material, etc. Generally, a relatively inert support is most desirable, since, under some conditions strong absorption of hydrocarbons by the support reduces good selectivity for discriminate oxidation between the feed and the product hydrocarbons. For example, activated chromatographic grades of alumina and silica gel absorb hydrocarbons very strongly and, unitl the materials are completely saturated with the hydrocarbons at the temperature of the reaction, no oxidation product is recovered. Fused alundum, some of the less porous and less absorbent silicas, alumina, and silica aluminas, and like material provide good support material. Amorphous silica gel precipitated by acidification from sodium silicate solution is an excellent support.

Equipment for the process of the invention may be suitable apparatus for either fluidized or fixed bed catalysis, and both types are known in many forms. One fluid bed reactor which is highly suitable for the process of the invention is shown in FIG. 1.

In conducting the selective oxidative cracking of the invention, it is highly important to maintain a close control on the temperature of the reaction so as to prevent the oxidation from proceeding beyond the desired dealkylation or" the aromatic hydrocarbons. The oxidative cracking of the invention proceeds rapidly and at a considerably lower reaction temperature than catalytic or thermal hydrocracking, and the contact times required for the catalytic and thermal hydrocracking are many times greater than for the oxidative cracking according to the invention. Uncontrolled oxidation, obviously, may proceed to complete oxidation with the production of carbon dioxide and Water, which is detrimental where the desired product is a parent aromatic homolog.

The reaction of the invention is conducted at closely controlled temperatures. A low air-to-feed ratio is necessary to obtain optimum production of the parent aromatics or other lower molecular weight from the higher homologs. Good selective oxidation conditions may be obtained in the region of from BOO-600 C. with a contact time of 0.05 to 1.0 seconds and, with an air feed ratio of 0.1 to 10 liters of air per gram of feed. Preferably, the temperature range is 400450 C. with up to about 10 weight percent of catalyst, and from about 5-25 weight percent ogygen.

The most convenient and practical gaseous oxidant is air. However, it has been demonstrated that the concentration of oxygen in the initial oxygen bearing stream is an important reactant variable which aiTects the selectivity of the oxidation. For certain applications the concentration of the gaseous oxidant in the gaseous phase may be varied from about 0.1% to nearly pure oxygen. Inert diluents, such as nitrogen, rare gases, carbon dioxide, carbon monoxide, and even water vapor may be added in controlled amounts to favor the oxidation selectivity. The pressures under which these reactions maybe effected range from about 0.1 to 1000 atmospheres with the practical range in 1 to 10 atmospheres.

In one form of the invention the oxidation process is conducted by passing a stream of vaporized hydrocarbon feed vertically upwardly through a column containing either a fixed bed or a fluid bed catalyst. The air-hydrocarbon feed rates are closely regulated by conventional means, and these are preheated to approximately the temperature of the desired reaction. The rate of feed of hydrocarbon and air is determined by the size of the column so as to limit the residence time of the hydrocarbon in the reactor to within a preferred contact time range depending on the particular catalyst.

Exemplary apparatus and oxidation results are shown in the appended drawings in which:

FIG. 1 is one form of a fluid bed oxidation catalyst apparatus;

FIG. 2 is a modified form of a fixed bed oxidation apparatus;

FIG. 3 is a graph showing the activity of three catalysts in producing naphthalene from l-methylnaphthalene;

FIG. 4 is a graph showing the effect of temperature on a feed with two mixtures of gas containing different amounts of oxygen; and

FIGS. 5 and 6 are charts showing isoyield and isoconversion contours of naphthalene from l-methylnaphthalene made by plotting the logs of Cadmium oxide concentration and oxygen concentration at 400 C. and 450 C., respectively.

In the form of the apparatus shown in FIG. 1 it consists of a tubular stainless steel reactor 1 having a lower inlet 2 for admission of air and hydrocarbon vapor, and an upper outlet 3 leading into a condensing train 4. The reactor provides a volume 5 in its upper portion which is of a larger cross-sectional dimension than the vertical column for catalyst disengaging, and in one form it has about a cross-section of six times the cross-sectional area of the lower zone 6, which is the reaction zone. This catalyst disengaging zone facilitates the settling of the catalyst from the vapor. A porous stainless steel disc 7 in the bottom of the reactor zone provides a catalyst support to prevent it from falling back through into the bottom of the reactor. This disc is sufficiently porous to provide only a small restriction in the fluid flow therethrough. Four pressure taps P P P and P are connected to pressure meters, for example, mercury-filled monometers (not shown) for measuring pressures in various parts of the reactor. In addition, by comparing the overpressures of the diiferent points in the reactor bed, it is possible to calculate the pressure drop across the catalyst bed and hence follow the extent of fiuidization. A porous stainless steel bayonet filler 9 covering the upper outlet, essentially prevents carry-over of catalyst into the condenser system. The reactor is suitably heated by wrapping electric heating units around the reactor and covering it with a glass cloth insulation. The heating elements are controlled by transformers (not shown), as is conventional practice. The temperatures in various sections of their reactor may be read on a temperature indicator (not shown) interconnected with the thermocouples TC TC TC and T0,.

In addition to the electric heating units wrapped around the reactor, another temperature controlling means is provided by means of a pressure condenser system 11 which includes a cold finger condenser portion 13 extending from an upper portion of the reactor chamber into the lower portion thereof. A thermostatic fluid is placed in the cold finger and the boiling point of this fluid is controlled by means of nitrogen pressure exerted on the system from the tank indicated.

Dry air from supply line 14 is divided into two

lines

15 and 16 which pass through rotameters 17 and 18, respectively, depending on whether the syringe pump 19 or carburetor 20 is used for hydrocarbon feed. The regulated air and hydrocarbon vapor streams are preheated in a preheater 21, which normally may be an electric element preheater. Air from the line 15 through the rotameter 17 enters the preheater 22 as primary air, and the hydrocarbon is introduced from the syringe pump 19 into the preheater along with the secondary air from line 16 through rotameter 18. The feed to the preheater is provided either as a liquid or a vapor, depending on whether the pump 19 or the carburetor 20 was used. Where the feeds are liquids or solids, they may be fed through a jacketed carburetor equipped with a heater.

The products of the reaction are passed from the reactor through outlet 3 into a glass U-tube air condenser system 23, bafile condensers 24, and generally through a series of three Dry Ice acetone bubbler traps 25 connected in series. The exhaust from the Dry Ice traps is measured in a wet test meter 26.

The catalyst used in the experiments set forth in Table 1 below is commercially available Davison 902 vanadia catalyst screened to between 100 +200 mesh, and which was preactivated in a ceramic tube for about twelve hours held at about 430 C. with a low flow of air thereacross to drive off the undesirable mineral acids and to further prevent corrosion of the oxidation apparatus. The composition of this Davison 902 catalyst is approximately vanadia, 33% potassium sulfate and 55% silica. This catalyst was conveniently diluted with inert Davison activated silica in approximately the same mesh size. The hydrocarbons oxidized in the test are set forth in Table 1 and these consisted of pure 2,3-dimethylnaphthalene and pure l-methylnaphthalene.

OPERATING PROCEDURE FOR FLUID CATALYST BEDS The fluid-like character of the fine catalyst permits ready addition of it to the reactor, as for example, from a weighed plastic squeeze bottle and it is readily removed by siphoning techniques without interrupting operating conditions. Prior to an actual run, an additional two hours activation time at about 375 and at about 500 6 liters of air per hour air flow in the reactor will bring the catalyst to temperature equilibrium with the reactor at about operating conditions. After such prescribed time and when the reactor temperature is essentially uni-' form in the range of 300400 C. a run may be commenced. The temperatures during these runs must be carefully controlled by adjustment of the transformers heating the reaction tube and by adjusting the nitrogen pressure on the cold finger condenser. The hydrocarbon feed is charged to the feed device and the run completed with a predetermined amount of hydrocarbon.

The total product trapped in the air condenser, the

baflie condenser, and the acetone traps from each run- Example I.-Oxidati0n of 2,3-dimethylnaphthalene to naphthalene About 219 ml. (212 g.) volume of Davison 902 fluid vanadia catalyst was activated in the apparatus at 430 C. for 2 hours at an air flow of 500 liters per hour. The fluid bed reactor was controlled at 375 C., and after temperature equilibrium was obtained, the syringe pump was heated and then filled with pure 2,3-dimethylnaphthalene. The air and feed rates were set to approxmately liters of air per gram of feed, and this gave about 0.4 second contact time of the vapor with the catalyst. run continued for a period of about 45 minutes. A total of 4.4 grams of 2,3-dimethylnaphthalene was fed through the reactor during this time at a rate of about 7.475 grams per hour. A total of 2.82 grams of condensable reaction product mixture was collected in the traps.

Calculated reaction conditions for the completed run were weight/hourly space velocity of 0.0343 kg. per liter-hour; gas/hourly space velocities of 4,560 liters per liter-hour; air-to-feed ratio of 133 liters of air per gram feed; and a contact time of 0.414 second.

Elution chromatography of about 1 gram sample of the reaction product mixture over chromatographic alurnina with eluting solvent mixture ranging from petro- 'leum ether through carbon tetrachloride, chloroform, and finally acetone yielded 23 fractions which were evaporated to dryness in tared flasks. A total of about 12.3% pure naphthalene was recovered from the product. This corresponds to a conversion of about 7.9 weight percent or 8.8 mole percent per pass.

The recovered naphthalene had a melting point of 80.l-80.7 C. A complete analysis of the material for naphthoquinone, phthalic anhydride, and maleic anhydride indicated a total hydrocarbon content of less than 12.9% by difference, and obviously most of this material is in the recovered naphthalene.

Examples 2 through 20.Oxidati0n of I-methylnaphthalene t'o naphthalene A total of 18 runs are summarized in Table I wherein pure l-methy-lnaphthalene was oxidized over commercial Davison 902 fluid vanadia, and the combinations of operation ranges include temperatures in the range of 300--351 0., contact times ranging from 0.101 to 0.761 second and an air-feed ratio of 4.71 to 20.20 liters of air per gram of feed. The yields of naphthalene range from 0 to about 8.1 weight percent conversion.

Operating the apparatus essentially is set forth in Example 1, and at the conditions set forth, the following table summarizes runs 3 through 20.

The air and feed flows were commenced and the TABLE I.-VAPOR-PHASE OXIDATION OF I-METHYL- NAPHIHALENE TO NAPHTHALENE OVER FLUID VANADIA Experimental Naphtha- Oonditions lene Re- Naphthalene Yield covcry Example Wt.per- T.( C.) t Sec Alf cent of Wt. per- Mole (l./g.) total cent a percent b I-LO.

A. PRELIMINARY 2 EXPERIMENTAL DESIGN B. ATTEMPT TO APPLY STEEPEST ASCENT TECHNIQUE WITH THREE VARIABLES $3 on about 5.5 ml. of preignited alundum. The bypass valves 40 are arranged to pass feed and air streams through the microreactor or to bypass the microreactor. The drying column 33 was a packed column of mixed ascarite and magnesium perchlorate to remove most of the carbon dioxide, water and organic acids which might interfere with the analysis of hydrocarbon from the two meter separation column 34 which contains silicon oil on firebrick. The hydrocarbons which emerged from the separating column are detected and assayed by detector cell 35 which was contained in a bridge circuit of an electronic chart recorder. The various feed mixtures were used for calibration by bypassing the microreactor to obtain the characteristic retention time, separations, and the peak areas for the naphthalene and a l-methylnaphthalene. The reaction product peak areas were used to determine the total hydrocarbon recovery and the amounts of naphthalene and l-methylnaphthalene in the recovered hydrocarbons. A pressure differential of about 4 p.s.i.g. across

air lines

36 and 37 was maintained at a fixed rotameter 32 air flow setting by means of fiow control valves 33 at the sample injection block 41. Hydrocarbon samples of 0.001 to 0.010 ml. were injected into the injection block 319 0.189 10. 03 0. 26 0. 0. 2s 1 3 0-191 11-64 0.36 0440 40 by means of calibrated microdippers. Tne metal ox- 00 20'20 25 ides investigated as catalysts for the production of naphthalcne accordin to this method are listed in Tabe II. 0. APPLICATION OF STEEPEST ASCENT TECHNIQUE g WITI-I ONLY TWO VARIABLES TABLE II.MTATERIALS INVESTIGATED 324 0.263 7.34 13.2 4. 63 5.14 Catalyst, Metallic Oxides Selectivity for Com- Production of N 325 0.260 11.94 28.0 0.31 0.34 bustion N/l-MN from l-MN 326 0. 735 7.49 27.2 2.90 3.21 325 0. 761 12. 55 17. 0 3. 50 3. 39 324 0. 14.9 0. 7s 7. 09 8. 03 8.96 325 0.154 5. 30 12. 32 4.10 1. 55 321 0. 154 0. 55 19. 22 4. 93 5. 47

1 a wti agrcent yield wt.naphthalenelwt. feedwt. l-mcthylnaphthalafole percent yield wt. percent yield X gmw l-mcthylnaphthalenc/ gmw naphthalene.

For Example 18 the l-methylnaphthalene, 71 1.6160, was prepared by fractional distillation and its isomeric purity confirmed by gas-liquid chromatographic analysis. The fluid bed reactor was brought to temperature equilibrium at about 325 C., and the preactivated fluid catalyst consisted of about 36.7 ml. (35.6 g.) of Davison 902 fluid vanadia catalyst at 100+200 mesh and about 100 ml. of Davison silica as a diluent. The l-methylnaphthalene was pumped into the apparatus at a rate of about 60 grams per hour and the air flow regulated at about 360 liters per hour based on standard temperature and pressure. The run was continued for a period of about one hour. The results are as set forth in Table I.

FIXED CATALYST BED A microreactor technique combined with gas-liquid chromatographic apparatus was used to investigate the selective catalytic oxidation reactivities of a large number of essentially pure reagent grade metal oxides for the production of naphthalene. Over 1400 runs were conducted with this technique and of the severaLfeeds studied, there were included a mixture of 25 weight percent naphthalene-75% substantially pure I-methylnaphthalene; a mixture of substantially pure 1,2- and 1,7-dimethylnaphthalenes; an aromatic extract from light catalytic cycle oil; toluene; l-naphthaldehyde; l-naphthyl alcohol; and 1- naphthoic acid.

One scheme of the apparatus for such a microreactor technique is shown in FIG. 2 wherein all the apparatus, with the exception of the air cylinder, was contained in two air baths of gas-liquid chromatographic apparatus controlled at 200 C. The microreactor tube 31 consisted of a stainless steel tube of about 0.63 centimeters ID. and about 40 cm. long. The tube 31 is covered by an electric heater 32 controlled by a temperature recorder-controller, as set forth above, controllable in the region of 300-600" C. A weighed samped of about 0.5 ml. of finely pulverized metal oxide was uniformly dispersed Notes:

a. Alundurn, an alumina/silica, gave little combustion at any temperature and was used as a support for testing the other oxides.

b. Silver oxide decomposes at 300 C. but naphthalene was produced over the decomposition product, which was probably the metal.

0. Antimony pentoxidc decomposes to 813204 at 380 C. and adsorbs all feed, yielding a new compound (unidentified) at 200 C. Complete combustion was noted at 300 C.

It is noted from Table I1 above that a substantial portion of the metal oxides investigated were found to be etiective in significant selective catalytic oxidation of the 25 weight percent naphthalene-75 weight percent 1- methylnaphthalene feed mixture at temperatures substantially below spontaneous ignition temperature of 559 C. as indicated by the plus marks of column 2 in Table II. Reagent grade vanadia was found to be more selective for the purpose than commercial Davison 902 vanadia catalyst, for example. Further, stable oxides of zinc,

silver, cadmium, indium and bismuth were found to be more selective than the reagent grade vanadia.

The results with the three most selective oxides of cadmium oxide, bismuth oxide, and indium oxide are summarized in FIG. 3. It is to be noted that if a catalyst manifested no selectivity, the composition of the recovered hydrocarbon remained substantially constant at the feed composition.

All of the metal oxides listed in Table III were also investigated for their production of naphthalene from pure l-methylnaphthalene. The production of naphthalene by these oxides is indicated in column 3 of Table III by the plus marks. The most effective catalysts for the production of naphthalene from l-methylnaphthalene were found to be cadmium oxide, bismuth oxide, indium oxide, and silver oxide. This production of naphthalene was achieved in the range of Soil-600 C. at about 0.1-2 seconds contact time at about 0.1 to 5 liters per gram air-to-feed ratios. Note that these conditions are considerably different from those of thermal or catalytic hydrocraclzing.

The carrier gas during one experiment was changed from air to nitrogen, and this confirmed the necessity of oxygen for the dealhylation since no naphthalene was produced with nitrogen as the carrier stream. Additional experiments were run to determine the eltect of oxygen concentration on the reaction, and this is shown in FIG. 4 wherein the upper line shows the conversion of l-methylnaphthalene to naphthalene in a oxygen carrier and in the lower curve of air (containing normal oxygen). The conversion is shown at different temperatures over 3.4-4.1 weight percent cadmiasilica (explained below).

A very reactive form of cadmium was prepared by the addition of cadmium nitrate to a solution of sodium silicate followed by coprecipitation with an acid. The recovered catalyst was very reactive and more selective than cadmium oxide reagent.

Examples 2139.Oxidati0n of J-methylnaphthalene over cadmium oxide l-methylnaphthalene was oxidized over cadmium oxide to produce directly naphthalene, and Example 31, detailed below, is typical of this set of runs. In this case, cadmium oxide was utilized for the selective vapor phase oxidation of 25 weight percent naphthalene-75 weight percent l-methylnaphthalene feed in air at a temperature range of 300-600 C.

In accordance with the procedures for the microreactor given above, the microreactor tube was filled with 0.5 ml. (0.400 g.) of finely pulverized cadmium oxide mechanically dispersed on about 5.5 ml. of -+60 mesh alundum support. The microreactor was brought into temperature equilibrium at about 525 C. by means of a thermocontroller.

A sample of 0.007 ml. of essentially pure l-methylnaphthalene was injected into air which was flowing at a rate of about 17.5 ml. per minute into the microreactor. The estimated reaction conditions are about 525 C.; O.22 seconds contact time; and 0.1 to 5 liters of air to grams of feed. The reaction product mixture gave a chromatograph curve consisting of carbon dioxide, water, and a naphthalene, in addition to unconsumed l-methylnaphthalene. An analysis of the curve shows that the recovered hydrocarbon amounted to about 58.5 weight percent of the l-methylnaphthalene feed, and of this 41.0 weight percent was naphthalene and 59.0 weight percent was l-methylnaphthalene. The conversion to naphthalene was about 24.0 weight percent per pass based on the total l-methylnaphthalene feed. The ultimate yield was found to be 36.6 weight percent or about 40.6 mole percent naphthalene based on the actual amount of the 1- methylnaphthalene consumed.

The results of the oxidation over this particular catalyst are summarized in Table IV. This table shows high sel0 lectivity of the cadmium oxide, and it was found by experimentation that cadmium oxide on an inert support exhibits better selectivity than the same amount of cadmium oxide and support disposed as a separate zone or a slug in the reactor.

TABLE III.SELECTIVE OXIDATION OF l-METHYLNAPH' THALENE TO NAPHTHALENE OVER OADMIUM OXIDE Conditions Hydrocarbon Recovery Example Feed Wt. Wt. 'lmf 0. Percent Percent Wt. Wt. N l-MN Percent Percent N l-MN N=Naphthalene l-MN =1Mcthy]naphthalene Other active forms of cadmium may be prepared by various chemical methods from decomposable salts or compounds of cadmium, e.g., cadmium nitrate, sulfite, carbonate, acetate and hydroxide. The more active and selective forms of the cadmium are those on various supports such as alundum, silica, silicon carbide and ceramic materials which may be prepared by evaporative or vacuum impre nation techniques with aqueous solutions of cadmium salts on the supports. Such catalysts, after being deposited on the support, are dried and treated at about 600 C. for two hours to decompose the cadmium compounds into the oxide.

One very active cadmium composition is produced when cadmium is precipitated in the presence of sodium silicate by the addition of an acid, after such precipitation which is recovered it is subsequently rinsed, and dried at 600 C. for about two hours. The catalyst recovered from this process is white in appearance. For purposes of this application such a catalyst is called cadmia-silica. The composition of the material obviously varies and is dependent on the amount of ingredients admixed together.

Example 40.Preparati0n o a cadmia-silica catalyst A solution of about 45 ml. of water glass in about 855 ml. of distilled water was stirred in a beaker. To this solution was added a solution of about 2.4 grams of cadmium nitrate in ml. of distilled water which was added drop-wise over a period of about ten minutes to produce a milky suspension. Four drops of phenophthalein were added and a third solution of aqueous nitric acid was added drop-wise with vigorous stirring over a period of about ten minutes. The pH of the final solution was about 6 as determined by indicator paper. The resultant mixture was then poured into centrifuge bottles and centrifuged, a supernatant solution decanted and the gel mixed and rinsed with portions of aqueous 10% ammonium nitrate and the resultant product dried for about 16 hours at 138 C. and then ignited at about 600 C. for two hours in a muifie furnace. The product was then screened to a -30+60. A sample of the catalyst was shown by emission spectroscopy to contain 3.4 weight percent cadmia. Diffraction X-ray analysis indicated the material to be of a rather amorphous structure.

1 1 SELECTIVE OXlDATION Following the procedure given above, the microreactor tube was filled with about 6.0 ml. (3.0 g.) of this catalyst. The operating procedure was similar to that described for Example 18 above. The results of the oxidation using this catalyst are summarized in FIG. 4 wherein the ultimate Weight percent of naphthalene yields were plotted for a temperature range of 400-600 C. In this case, two carrier gas streams were employed, as pointed out above, one at 20% oxygen (air) and the other at 10% oxygen, which is air diluted with an equal volume of nitrogen. The results obtained with these different oxidation mixtures are presented on curves A and B of FIG. 4. It is noted that better yields are obtained at the lower temperatures and with a lower oxygen concentration over this particular catalyst.

Examples 41-48.-Selective oxidation of various feeds Following the procedure of Example 18, the microreactor-gas-liquid chromatographic technique was used to investigate the selective oxidation of several feeds other than the monomethylnaphthalene. In Examples 41 and 42, the production of monom-ethylene and naphthalene are demonstrated by the selective oxidation of 1,7- and 1,2- dimethylnaphthalene (1,7-DMN and 1,2-DMN). Example 43 shows the oxidation of an aromatic extract of light catalytic cycle oil from a catalytic cycle oil in the boiling range or" 490-525 C. This feed, which contains various alkylaromatic compositions, Was selectively oxidized to produce monomethylnaphthalene and naphthalene. Example 44 shows the selective oxidation of 1- naphthaldehyde to produce naphthalene, and in Examples 45 and 46 decalin and tetralin, respectively, were oxidized to produce naphthalene. Example 47 shows the results of oxidizing toluene to produce benzene, and Example 48 shows the production of benzene and toluene from the oxidation of xylene. The catalyst for Examples 41a, 42a and 43 was cadmium oxide dispersed on alundum, and for the remainder of the examples in the table, the catalyst is cadmium coprecipitated with silica, as described above. The results of these tests are shown in Table V, given below.

TABLE IV.-OXIDA'lION OF VARIOUS FEEDS TO I-IYDROCARBONS Example Conditions Reaction Products Feed Oxidant 1,7-DMN.

N, MMN, DMN. N, MMN, DMN. Naphthalene.

Do. Benzene. Benzene, toluene.

N=naphthalcne. MMN=monon1ethylnaphthalcne. DMN dimethylnaphthalenc.

Examples 4966.C0l1versi0n f l-methylnaphtlmlene where a is the weight percent yield of naphthalene, x is the log (Wt. percent CdO) on the catalyst, y is the log (Wt. percent 0 in the carrier gas stream, and A,

12 B, C, D, E,'and F are empirical constants. The conversion data at each temperature were fitted to a seconddegree polynomial equation of the general form, (2) 6:1 1+Bx+Cy+Dx +Ey -i-Fxy where B is the weight percent conversion to naphthalene, x is the log (wt. percent CdO) on the catalyst, y is the log (wt. percent 0 in the carrier gas stream, and A, B, C, D, E and F are empirical constants. The yield data at 400 C., with the exception of Example 50, were titted to a particular polynomial to :5 .4 wt. percent; the conversion data at 400 C., with the exception of Ex ample 50, were fitted to a particular polynomial to :31 wt. percent. All of the yield data at 450 C. were fitted to a particular polynomial to :3.8 wt. percent. All of the conversion data at 450 C. were fitted to a particular polynomial to i2.8 Wt. percent.

These experimental results at 40 C., and 450 C. are graphically represented in FIG. 5 and FIG. 6, respectively. The data points are indicated by the circled numbers and refer to the key in Table V. The isoconversion and isoyicld contours were obtained from graphic solutions of the empirical polynomial expression which obtained from least square fits of each group of data. On inspection of the regions of overlap for the isoconversion and isoyield contours, it is apparent that the more favorable reaction conditions at 400 C. are over about 2-5 wt. percent CdO, with about 525 wt. percent 0 in the carrier gas stream, yielding about from 30 to 62:5 wt. percent naphthalene at conversions of about 2025:L3 wt. percent, and at 450 C., over 2-8 wt. percent CdO, with about 5-25 wt. percent 0 in the carrier gas stream, yielding about from 32 to 58i4 wt. percent naphthalene at conversions of about 20-25i3 Wt. percent.

It is thus apparent that selective oxidative dealkylation of higher homologous alkyl substituted hydrocarbons is a feasible and economical process for producing lower member of such a series. It is, furthermore, a valuable process for recovering parent aromatic hydrocarbons from a mixture of alkyl substituted compositions of such parent compounds. While the invention has been described by reference to specific examples, there is no intent to limit the spirit or scope of the invention to the details so set forth, except as defined in the following claims.

TABLE V.--EFFECTS OF REACTION VARIABLES ON THE SELECTIVE OXIDATION OF l-METHYLNAPI-ITHALENE TO NAIHTHALENE Conditions Results Key No. Example on Figs. Wt. Per- Wt. Per- No. 5 & 6 Temp, Wt. Wt. cent Con cent Yield 0. Percent Percent version to of Naphtha- O2 CdO Naphthalone lens We claim:

1. A selective dealkylation oxidation process for the production of lower molecular weight aromatic hydrocarbons from higher molecular weight alkyl substituted aromatic hydrocarbons which comprises contacting for a limited time a vaporous mixture of such alkyl substituted aromatic hydrocarbons and a gas containing from 1 to 50% of oxygen by weight with a fluidized bed of an oxidation catalyst selected from the group consisting of inorganic insoluble salts of cadmium and insoluble oxides of silver, zinc, cadmium, indium, and bismuth, in a temperature range of from 300-600 C. and at a gas-tofeed ratio of from 0.5 to 10 liters per gram whereby to limit the residence of the hydrocarbon over the catalyst at from .05 to seconds.

2. A- process according to claim 1 in which the oxidation catalyst is an insoluble, inorganic cadmium salt.

3. A process according to claim 1 in which the oxidation catalyst is cadmium silicate.

4. A process according to claim 1 in which the oxidation catalyst is cadmium oxide.

5. A process according to claim 1 in which the oxidation catalyst is a co-precipitated product of a soluble cadmium salt and a water glass solution.

6. A process according to claim 1 in which the oxidation catalyst is bismuth oxide.

7. A process according to claim 1 in which the oxidation catalyst is indium oxide.

8. A process for the production of naphthalene from l-methylnaphthalene which comprises contacting for a limited time a vaporous mixture of such l-methylnaphthalene and a gas containing about 5-25 weight percent of oxygen with an oxidation catalyst selected from the group consisting of inorganic in soluble salts of cadmium and insoluble oxides of silver, zinc, cadmium, indium and bismuth, at temperatures of between 400-450" C., the ratio of gas to feed being on the order of from 0.5 to liters per gram whereby to provide a residence time of the aromatic hydrocarbon with the oxidation catalyst of from 0.1 to 2 seconds, and separating the produced naphthalene from the mixture.

9. A process according to claim 8 in which the oxidation catalyst is cadmium oxide.

10. A selective oxidation process for the production of lower alkyl homologs of benzene and benzene from higher molecular weight alkyl substituted benzenes which comprises contacting for a limited time a vaporous mixture of such alkyl substituted benzenes and a gas containing from 1 to oxygen by weight with a fluidized bed of an oxidation catalyst selected from the group consisting of inorganic insoluble salts of cadmium and insoluble oxides of silver, zinc, cadmium, indium, and bismuth in a temperature range of from 300-600 C. and at a gasto-teed ratio of from 0.5 to 10 liters per gram whereby to limit the residence time of the hydrocarbon over the catalyst of from 0.05 to 5 seconds.

References Cited in the file of this patent UNITED STATES PATENTS 2,301,735 Melaven et a1 Nov. 10, 1942 2,343,450 Free et al. Mar. 7, 1944 2,351,793 Voorhees June 20, 1944 2,370,541 James Feb. 27, 1945 2,470,411 Corner May 17, 1949 2,474,002 Levine et al June 21, 1949 2,565,627 Pryor Aug. 28, 1951 2,810,764 Steadman et a1. Oct. 22, 1957 FOREIGN PATENTS 159,508 Great Britain Aug. 28, 1922 649,999 Great Britain Feb. 7, 1951 OTHER REFERENCES Chemical Abstracts (I), vol. 41, page 5465f (1947). Chemical Abstracts (11), vol. 43, page 7916;; (1947).

Claims

1. A SELECTIVE DEALKYLATION OXIDATION PROCESS FOR THE PRODUCTION OF LOWER MOLECULAR WEIGHT AROMATIC HYDROCARBONS FROM HIGHER MOLECULAR WEIGHT ALKYL SUBSTITUTED AROMATIC HYDROCARBONS WHICH COMPRISES CONTACTING FOR A LIMITED TIME A VAPOROUS MIXTURE OF SUCH ALKYL SUBSTITUTED AROMATIC HYDROCARBONS AND A GAS CONTAINING FROM 1 TO 50% OF OXYGEN BY WEIGHT WITH A FLUIDIZED BED OF AN OXIDATION CATALYST SELECTED FROM THE GROUP CONSISTING OF INORGANIC INSOLUBLE SALTS OF CADMIUM AND INSOLUBLE OXIDES OF SILVER, ZINC, CADMIUM, INDIUM, AND BISMUTH, IN A TEMPERATURE RANGE OF FROM 500-600*C. AND AT A GAS-TOFEED RATIO OF FROM 0.5 TO 10 LITERS PER GRAM WHEREBY TO LIMIT THE RESIDENCE OF THE HYDROCARBON OVER THE CATALYST AT FROM .05 TO 5 SECONDS.