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Publication numberUS3926781 A
Publication typeGrant
Publication dateDec 16, 1975
Filing dateOct 9, 1973
Priority dateOct 9, 1973
Publication numberUS 3926781 A, US 3926781A, US-A-3926781, US3926781 A, US3926781A
InventorsLaird H Gale
Original AssigneeShell Oil Co
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Catalytic cracking of paraffinic naphtha
US 3926781 A
Abstract  available in
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Claims  available in
Description  (OCR text may contain errors)

United States Patent Gale Dec. 16, 1975 CATALYTIC CRACKING F PARAFFINIC 2,096,769 /1937 Tropsch 2,889,268 6/1959 Dinwiddie et a1 NAPHTHA 3,310,597 3/1967 Goble et a1. 260/683 Inventor: Laird Gale, ust n, Tex. 3,770,616 10/1973 Kominami et a1. 208/138 Assignee: She Oil p y, Houston Tex 3,772,184 10/1973 Bertolacinl et a1. 208/65 [22] Filed: Oct. 9, 1973 Primary Examiner-Delbert E. Gantz [211 App! 404541 Assistant Examiner-G. E. Schmitkons [52] US. Cl. 208/117; 208/122; 208/135; 208/141; 252/442; 252/463; 260/673.5;

260/677 R [57] ABSTRACT [51] Int. C1. ..........................Cl0G 11/08; B01] 27/12; I

BOlJ 23/08; C07C 11/02 A gallia-alumina or fluorided gallia-alumina catalyst is Field of Search 208/115, 1 6, 1 used for cracking paraffin-containing hydrocarbon dis- 208/1 260/6735 tillate feedstocks to produce light olefins and highly aromatic gasoline. [56] References Cited UNITED STATES PATENTS a8 (Ilaims, 1 Drawing Figure 1,935,177 11/1933 Connally at al 252/456 CATALYST B E it Q 16 6' CATALYST A 0 I I I TIME I MINUTES U.S. Patent Dec. 16, 1975 CATALYST B CATALYST A I T/ME, MINUTES BACKGROUND OF THE INVENTION This invention relates to the catalytic cracking of hydrocarbons to produce products boiling below the boiling range of the hydrocarbons cracked. In particular it relates to the catalytic pyrolysis of paraffins to produce light gas and aromatics.

A fluorided alumina catalyst has been used to catalytically crack refinery feedstocks. A process designed to produce normally-gaseous olefins having a high propylene content is described in U.S. Pat. No. 3,310,597.

Substantial aromatization activity results in cracking hydrocarbons over alumina and fluorided alumina, but the yields are poor. A catalytic pyrolysis process which has a high conversion of hydrocarbons to lower boiling products containing a high yield of light olefins and gasoline boiling range aromatics would be of great value in view of the increasing demand for these products. Accordingly, it is an object of this invention to provide a cracking process which utilizes the intrinsic cyclization-aromatization activity of a modified alumina catalyst to accomplish such an improved product distribution. In particular it is an object of this invention to convert paraffinic hydrocarbon feedstocks to lower molecular weight aromatics by a process which can be described as "dehydrocracking-aromatization" (DCA).

SUMMARY OF THE INVENTION A catalytic cracking process which comprises contacting a paraffin-containing hydrocarbon distillate feedstock boiling substantially in the range C,,-450C at cracking conditions with a catalyst comprising a major proportion of alumina combined with about 2 to 40% wt gallia to obtain a product containing substantial amounts of normally gaseous olefins and gasoline boiling range aromatics. The catalyst may optionally contain from about 1 to 5% wt fluoride.

DESCRIPTION OF DRAWINGS The FIGURE shows the effect of catalyst age in a process for crackiing cumene to alpha-methylstyrene with alumina and gallia-modified alumina catalysts.

DETAILED DESCRIPTION The present invention is concerned with a catalytic cracking process utilizing a catalyst which is particularly selective in producinglight normally gaseous olefins and gasoline boiling range aromatics from paraffincontaining distillate hydrocarbon feedstocks boiling substantially in the range C 450C Such a conversion process, which operates, e. g., by converting a paraffinic feed to a lower molecular weight aromatic product, can be called a dehydroeracking-aromatization (DCA) process.

Since preliminary experiments established that pure alumina exhibits an interesting cyclization-aromatization activity when cracking normal paraffins, the properties of alumina were studied to see how they affected this activity. Two properties were considered most important. The intrinsic acidity of the alumina as described by Pines and coworkers in .I. Am. Chem. Soc., 82, 2471 (1960); and its crystalline modification, i.e.. eta, gamma, chi. etc.

The most acidic aluminas are prepared from hydrolysis of aluminum isopropoxide while the least acidic aluminas are prepared from sodium or potassium aluminate. Samples of alumina were prepared by each method and it was determined that both aluminas had the eta crystalline structure. These aluminas were used to crack pure n-octane at a temperature of 580C, atmospheric pressure and 1.6 WHSV. These tests showed that the selectivity to aromatics was greater with the nonacidic alumina prepared from sodium aluminate. Additional n-octane cracking tests with a gamma and chi alumina showed that the eta form has the highest selectivity to aromatics. Accordingly, the eta form is preferred and has been used in the galliaalumina catalysts of the invention.

The n-octane cracking studies showed that an alumina with a high alkali metal content (3.3% wt sodium) had virtually no cyclization-aromatization activity. Accordingly, a low-sodium, eta alumina was used as a standard in determining the effect of reaction variables on cyclization-aromatization activity. It was concluded that effective catalysts of the invention should have an alkali metal content of about 0.2% wt or less.

The gallia-alumina catalysts of the invention may be prepared in various ways, e.g., by impregnation of the alumina with gallia or by coprecipitation of the gallia and alumina. The latter method is illustrated in Example II below. The catalyst surface area should fall in the range of about 50 to about 300 m /g.

The catalysts of the invention contain a major proportion of alumina, preferably eta alumina, and a minor proportion of gallia, e.g., from about 1 to about 40% wt gallia, with from about 3 to about 10% wt gallia being a preferred composition range.

In addition to gallia and alumina the catlysts of the invention may contain from about 1 to about 5% wt fluoride to enhance acid cracking activity of the catalyst. A gallia-alumina catalyst of the invention containing 1.5% wt fluoride was found to be particularly effective in cracking a highly paraffinic hydrocracker recycle oil.

The operating temperature range for the process is from about 500 to about 625C. Conversion increases strongly with temperature while selectivity to aromatics reaches a maximum in the range of 560 to 590C. The proportion of light gases in the product increases rapidly with increasing temperature but the quality of the gases is lower, i.e., less olefinic. Furthermore, the deposition of coke on the catalyst increases with increasing temperature.

Suitable operating pressures for the process range from atmospheric (0 psig) to about 50 psig. Preferably the pressure ranges from O to about 15 psig.

Suitable weight hourly space velocities (WHSV) for the process range from about 0.5 to about 6. Preferably the WHSV will be from about 1 to 2. The aromatic content of the gasoline fraction increases markedly with decreasing WHSV, as does the proportion of light gases. In addition the light gases become increasingly saturated and thus less valuable. In general it may be said that reaction conditions which favor the highest aromatic content of the gasoline product yield large gas fractions of relatively low olefin content.

. Suitable hydrocarbon feedstocks for the process include paraffin-containing distillates boiling substantially in the range C,,-450C. Preferably the range will be from about C -450C. Since the process operates by selectively dehydrocracking-aromatization (DCA) of paraffins it is necessary that the feedstock contain a substantial proportion of paraffins. Preferably the feedstock will contain from about 40 to 100% v paraffins. The process is particularly effective in the DCA of pure paraffins such as n-octane and n-dodecane. However, the catalysts of the invention are also useful for the selective DCA of actual refinery feedstocks having paraffin contents falling within the preferred range.

When processing such refinery feedstocks it is generally preferred that fluoride be added to the catalyst to enhance its selectivity to aromatics in the gasoline boiling range.

The catalysts of the invention will generally be applied in a fluid bed process where frequent catalyst regenerations are required. A fixed-bed process can be used where infrequent regenerations are required.

It has been observed that catalyst aging has three general effects on the process of the invention:

1. conversion level increases slightly with catalyst age;

2. selectivity to aromatics increases initially (-0.5 hour sample vs. 0.5-1.5 hour sample) and then declines somewhat; and

3. the product distribution changes significantly with time. As the catalyst ages the degree of skeletal isomerization of the C -olefin fraction decreases and the C aromatic product distribution shifts toward o-xylene and ethylbenzene.

Because of the decline in DCA activity with time an alumina catalyst was processed for a 5 hour period with n-dodecane at 0.8 WHSV (to promote rapid coke deposition). Total products were collected and analyzed. The catalyst was then regenerated by combustion of the coke with air at 580C. The catalyst was then used to process n-dodecane for a 3-hour period, during which total products were again collected and analyzed. A comparison of these test results showed that regeneration had essentially no effect on conversion level and only a modest (5- l 0%) decrease in selectivity to aromatics. This test suggests that the gallia-alumina catalysts of the invention are amenable to regeneration.

The'invention will now be further illustrated by the following examples.

EXAMPLE 1 An eta alumina Catalyst A was prepared as a basis for comparison by a method similar to that given by Pines and Haag, J. Am. Chem. Soc. 82, 2471 (1960).

' The preparation method was as follows: 123 g sodium aluminate was dissolved in 3 liter distilled water. Carbon dioxide was bubbled in until no more precipitate was formed. The resulting aluminum hydroxide was collected by filtration and washed repeatedly with reslurrying to remove sodium ions. The product was dried at l 10C for 2 days. The aluminum hydroxide was converted to eta alumina by calcining in air at 550C for 16 hours. The finished alumina catalyst had a surface area of about 210 m /g and a sodium content of about 0.17% wt.

EXAMPLE 11 A gallia-alumina Catalyst B of the invention was prepared as follows: One mole of aluminum chloride was dissolved in 3 liters of water. The pH was adjusted to 7.0 with 6N ammonium hydroxide and a solution of 0.045 mole gallium chloride dissolved in 500 ml water was added. After mixing, the pH was further adjusted to 9.5 with ammonium hydroxide. The resulting gel was aged overnight. The gelled catalyst was washed six times with 0.1 N NH OH and dried for 4 days at C. The resulting solid was crushed and meshed to the desired size. Finally, the catalyst was calcined in air for 16 hours at 550C before use. The finished gal1iaalumina Catalyst B had a gallia content of 7.64% wt, a surface area of about 200 m /g and a sodium content of less than 0.1% wt.

EXAMPLE lll Catalysts A and B were used in a dehydrocrackingaromatization (DCA) process to convert n-paraffin hydrocarbons to aromatics. The feedstock for this example was pure n-dodecane. The feedstock was delivered by a syringe pump to an all-glass reaction system. The reactor was a %inch OD 17-inch long Vycor tube. which had a catalyst bed volume of 19 cc and which was heated by a three-section Lindberg l-leviduty Type 705 electric furnace.

A typical catalyst charge consisted of 2 g of catalyst (30-45 mesh) dispersed in 10 g ofquartz chips. A preheat section of the tube was filled with quartz chips. Liquid reaction products from the process were condensed in a water cooled condenser and collected in an efficient glass trap in an ice bath, while gaseous reaction products were taken out of the system through a wet test meter. Representative gas samples were collected in glass sampling vessels.

Gaseous reaction products were analyzed by mass spectrometry. The liquid products were analyzed by gas-liquid chromatography (GLC) using a 14inch OD x 23.5-foot SF-96/Chromosorb W (acid washed, Hexamethyldisiloxane treated) column held at 30C for 9 minutes followed by programmed heating from 30 to 250C at 2/minute. Total analysis time is about 2 hours. Peak identification was accomplished by combining retention time data developed from known compounds and GLC and mass spectrometry analyses. For samples requiring resolution of the C3 aromatics a Ainch OD X 20-foot Bentone 34/diisodecylphthalate column was used. Total coke yields were obtained by a combustion technique.

A Fortran V computer program was written to perform the laborious calculations required to combine the gas and liquid product analyses and coke yield into one overall product distribution. Operating conditions and test results from these comparative DCA processes are shown in Table I.

Table 1 Temperature: 580C Pressure: Atmospheric Table l-continued Table 3 Temperature: 580C SRHGO HRO Pressure: Atmos heric WHSV; 46 Gravity API 32.3 44.2 Time: 2.0 hr. Bp Range (GLC). 72w Expt. No. l I 2 Start 82C 0.1 Catalyst A B 82 160C 0.8 2.6 C-l2 Aromatics 8.5 2.8 19.3 8 9 160 199C 1.3 23.8 199C 216C 1.4 28.3 "'1 mole gallium/21) moles aluminum. 216 271C 10.6 31.8 Average of two experiments. 271 86.0 13.5 "Normalized to 100%. 10 Average Molecular Weight 294 I78 'Moles product/101) moles n-dodecane reacted. Composition. 7rw "C ubstituted benzenes naphthalene Paraft'ms 3 l .3 59 "C -substituted benzenes methylnapthalenes Naphthenes 50.3 32 "C,;-Sl.lblilllld benzenes dimethyland ethylnaphthalenes. Aromatics 1 9 U.V. Aromatics, mM/100 g Mono- 56.5 31.3 Table I shows that the incorporation of gallia into alumina (Catalyst B) increases the selectivity to aromat- :55

Total 62.2 32.3

ics. This is accomplished apparently by increasing the dehydrogenation activity of the catalyst. Increased dehydrogenation activity should increase the contribution to aromatics formation from dehydrogenation to trienes followed by thermal cyclization. Evidence for enhanced dehydrogenation activity is shown by increased yields of C12 aromatics from n-dodecane.

EXAMPLE IV The enhanced dehydrogenation activity of the galliaalumina Catalyst B was further demonstrated by cracking cumene in a process similar to that described in Example 111. Operating conditions and test results are shown in Table 2 and the FIGURE. The FIGURE demonstrates that the yield (wt. plotted as GLC area) of the dehydrogenative product from cumene, alphamethylstyrene, is much higher from cumene cracking on gallia-alumina Catalyst B compared to alumina Catalyst A. The initial absolute yield of benzene from the Catalyst B is nearly comparable to that from pure alumina indicating comparable concentrations of strong acid sites.

EXAMPLE V Catalysts A and B and a commercial zeolite cracking Catalyst C (Davison DZS were used in a DCA process to crack two refinery feedstocks. The feedstocks used for these experiments were a hydrotreated straight run heavy gas oil (SRHGO) and a second stage hydrocracked recycle oil (l-IRO) with properties as shown in Table 3.

Table 2 A simple yet accurate test procedure using relatively small quantities of catalyst and feed was developed to compare catalysts. The apparatus consisted of a fixedbed microcata-lytic system utilizing the all-glass microflow reactor as described in Example 111. Detailed product yield structures were obtained by analyzing product placed in the Vycor glass reactor. The catalyst bed was heated to 580C with N purge over a 30 min period and held at 580C for 1 hr with flowing N The N was then replaced by liquid feed at 7.5 g/hr. Liquid product was collected for a I hr period. Several representative gas samples were collected and the total volume of gaseous products was measured using a wet test meter. Following product collection, the catalyst bed was purged for 1 hr with N at 580C. Coke analyses were made by contacting the catalysts with air and trapping the resulting CO in aqueous sodium hydroxide. A

heated CuO bed insured completed combustion of CO to CO A titration procedure described by Pines and Csiscery, J. of Catalysis 1, 313 (1962), was used to determine the carbonate concentration in the aqueous sodium hydroxide solution. Typical material balances of 97% or better were obtained using these procedures. The gas samples were analyzed by mass spectrometry while the liquid products were analyzed by temperature Feed: cumene/Helium N1 Temperature: 580C Pressure: Atmospheric WHSV: 5.8 Expt. No. 3 4 Catalyst A" B"' Time. hr 0-1 1-2 01 1-2 Product Distribution. TLP 7rw *b) %w b) 7rw *b) kw *b) Benzene 5.92 25.0 3.53 .3 5.39 13.9 1.08 4.6 Toluene 0.39 1.4 0.20 .88 0.59 1.3 Ethylbenzene 3.28 10.0 2.41 .0 1.96 3.7 0.45 1.4 Styrene 3.36 10.5 1.91 4 2.20 4.2 0.63 2.0 Cumene 67.67 72.54 44.56 64.99 n-Propylbenzene 2.67 7.3 2.01 .7 2.05 3.4 a-Methylstyrene" I 1.33 30.9 13.0 .1 36.60 62.2 30.87 86.5 trans-B-Methylstyrene 5.41 14.9 4.42 3 6.65 l 1.3 1.98 5.5 Conversion 35.3 5.0 29.5 i 4.3 57.3 35.8

"'Includes some cis-B-merhylstyrene. '"Moles product/100 moles cumene reacted. Surface area. 247 sqm/g.

"'Average of two experiments.

"1 mole gallium/20 moles aluminum programmed GLC. The results of the gas, liquid and coke analyses were combined into a single overall product yield structure by a Fortran V computer program.

High boiling products and unconverted components in the feed boiling range could not be determine'd'directly by GLC. These products were determined indirectly by adding an internal marker (%w methylcyclohexane) and then relating each observed peak area to the known amount of marker. The difference between the sum of the %w for observed peaks and 100% .is the amount of undetected higher materials. The accuracy of the internal marker technique was checked for sev-. eral products by submitting these samples for GLC boiling point analysis which is capable of detecting hydrocarbons up to C The quantities of product boiling 271C determined by the GLC boiling point analysis and the internal marker technique were in good agreement.

The analysis of the products from cracking the hydrocracker recycle oil were complicated by the fact that the-feed initial boiling point (23%w 160-199C) overlapped the aromatic products in part of the heavy gasoline range. The yields of benzene, toluenes, ethylbenzene and xylenes could be determined directly from the GLC analysis of the TLP. The yields of higher alkyl aromatics C -C carbon number) were determined fromhigh resolution mass spectral analysis.

The results of cracking tests on the hydrotreated SRHGO feedstock comparing gallia-alumina Catalyst B of the invention with an eta alumina Catalyst A and a commercial type zeolite Catalyst C, (Davison DZ-5) are shown in Table 4. Zeolite cracking catalysts are noted for their high conversion yields and selectivity to an aromatic gasoline fraction.

Table 4 Feed: Hydrotreated SRHGO Pressure: Atmospheric Table 4-continued Feed: Hydrotreated SRHGO Pressure: Atmospheric "1 mole gallium/20 moles aluminum 1 hr. reaction time "'2 hr. reaction time Thesedata show that there is no significant difference'i'n the yields of heavy gasoline and the aromatic contents of the heavy gasoline fraction for the eta alumina Catalyst A and the zeolite Catalyst C at 580C and atmospheric pressure. However, the hydrogen yield is much higher for Catalyst A indicating higher overall aromatization activity for the alumina. The bulk of this additional aromatization apparently yields polycondensed aromatics which end up in the feed boiling range or as coke. The alumina promoted cracking (Catalyst A) yields a poor light olefin distribution. Compared to zeolite catalytic cracking (Catalyst C), alumina cracking produced considerably'more C and C product relative to C and C Also the iso/normal ratio of paraffins is much lower for thelight gas gasoline fraction (C /C from alumina cracking.

The gallia-alumina Catalyst B of the invention, which exhibited enhanced cyclication/aromatization activity with pure n-paraffins, was less active than the eta aluggg g v A I g g mina Catalyst A with this SRHGO feedstock. This poor Temp-emurefc v580 560 580 580 actiyity resulted from rapid deactivation by excessive WHSV 1.5" .0" coking. Apparently, the intrinsic cyclization/aromatiza- Product tion activity of Catalyst B did not result in a higher yield Diisirilputionfiw 2 2 2 6 l 4 0 6 of gasoline boiling range aromatics because of the low y rogen 7 Methane 59 Z8 42 36 paraffin content (31.3%;v) of thlS feedstock. Ethane 5.0 1.8 4.0 3.0 The results of cracking tests on the second stage hydrocracker recycle oil (HRO) comparing a galliapropylene 13 5,0 72 alumina catalyst of the invention with an eta alumina, gu a ne g2 and with fluorided versions of these catalysts, as well as U V enes 6 HC Gas 285 13] 31,8 with the same commercial zeolite cracking catalyst (Davison DZ-5) are shown in Table 5.

Table 5 Temperature: 580C Pressure: Atmospheric WHSV: 1.6 Time: 1.0 hr. Expt. NO. 9 10 11 12 13 Catalyst A A+ 'n B+" C Product Distribution. 7cw

Hydrogen 2.1 1.5 3.2 2.8 0.5 Methane 3.3 4.2 3.0 3.2 2.6 Ethane 2.6 1.7 2.2 2.0 1.8 Ethylene 2.3 2.9 2.1 4.1 2.0 Propane 1.8 3.2 1.6 1.8 2.4 Propylene 4.0 10.8 2.9 6.9 8.7 Butane 1.3 4.0 1.4 3.6 5.4 Butylenes 4.2 10.4 4.1 7.2 9.2 Sum HC GAS 20.5 37.2 17.3 28.8 32.1 Light Gasoline (C /C) 7 2 s s 5.6 7 2 13 5 Heavy Gasoline (CT/200C) 32.4 20.9 36.0 29.6 26.6

Table -continued Temperature: 580C Pressure: Atmospheric "'7.647( wt gallia Cracking of the l-IRO on a pure eta alumina (Catalyst A) yields 33% more gasoline range aromatics than cracking on the zeolite (Catalyst C). Cracking of this practical feedstock over the gallia-alumina (Catalyst B) gave further improvement in the yield of gasoline aromatics.

In order to improve the light olefin product distribution from alumina cracking, the acid cracking activity was increased by incorporating fluoride into both pure alumina and gallia-alumina catalyst. The data in Table 5 show that cracking the recycle oil on fluorided alumina does give improvements in the yields of propylene and butylenes; however, the yield of gasoline range aromatics is adversely affected. On the other hand, the 1.5%w fluorided gallia-alumina catalyst results in substantial increases in both propylene and butylene yields while the aromatics yield is only slightly reduced compared to pure gallia-alumina. The yield of gasoline range aromatics is 64% greater for Catalyst B 1.5% F than that of Catalyst C. The gallia-alumina catalyst B 1.5% F does, however, have a higher coke yield than that for Catalyst C, i.e., 10.2% versus 2.9%.

What is claimed is:

l. A catalyst cracking process which comprises contacting a hydrocarbon feedstock containing 40 to 100%v paraffms and boiling substantially in the range C -45OC, at cracking conditions, with a catalyst consisting essentially of alumina containing from about 1 to 40% wt. gallia, and recovering a product having a major portion of hydrocarbons boiling below the boiling range of the feedstock and containing substantial amounts of normally gaseous olefins and gasoline boiling range aromatics.

2. The process of claim 1 wherein the cracking conditions include a temperature of about 550 to 625C, an operating pressure of 0 to about 50 psig, and a weight hourly space velocity of about 0.5 to 6.

3. The process of claim 2 wherein the catalyst consists essentially of an eta alumina containing about 3 to 10% wt. gallia.

4. The process of claim 2 wherein the catalyst has a surface area from about 50 to about 300 m /g.

5. The process of claim 4 wherein the catalyst has a sodium content of about 0.2% wt or less.

6. A catalytic cracking process which comprises contacting a hydrocarbon feedstock containing 40 to %v paraffins and boiling substantially in the range C -450C, at cracking conditions, with a catalyst consisting essentially of alumina containing from about 1 to 40% wt. gallia and from about 1 to 5% wt fluoride, and recovering a product having a major portion of hydrocarbons boiling below the boiling range of the feedstock and containing substantial amounts of normally gaseous olefins and gasoline boiling range aromatics.

7. The process of claim 6 wherein the alumina is eta-alumina, and the alumina contains from about 3 to 10% wt. gallia.

8. The process of claim 6 wherein the hydrocarbon feedstock contains from about 40 to 100%v paraffms and boils substantially in the range C -450C, and the catalyst has a surface area from about 50 to about 300 m /g and contains less than about 0.2% wt alkali metal. =l

Patent Citations
Cited PatentFiling datePublication dateApplicantTitle
US1935177 *Dec 9, 1930Nov 14, 1933Silica Gel CorpProcess of preparing catalytic masses
US2096769 *Dec 24, 1934Oct 26, 1937Universal Oil Prod CoTreatment of hydrocarbon oil
US2889268 *Mar 6, 1953Jun 2, 1959Exxon Research Engineering CoMethod of converting hydrocarbons using a heat stable gamma alumina
US3310597 *Jun 13, 1962Mar 21, 1967British Petroleum CoCatalytic cracking
US3770616 *May 5, 1971Nov 6, 1973Asahi Chemical IndMethod of reforming hydrocarbons
US3772184 *Dec 17, 1971Nov 13, 1973Standard Oil CoReforming petroleum hydrocarbons with catalysts promoted with gallium and rhenium
Referenced by
Citing PatentFiling datePublication dateApplicantTitle
US4056575 *Jul 12, 1976Nov 1, 1977The British Petroleum Company LimitedChemical process making aromatic hydrocarbons over gallium catalyst
US4056576 *Jul 12, 1976Nov 1, 1977The British Petroleum Company LimitedChemical process over gallium catalyst converting saturated hydrocarbons to olefins
US4157356 *Dec 2, 1977Jun 5, 1979The British Petroleum Company LimitedProcess for aromatizing C3 -C8 hydrocarbon feedstocks using a gallium containing catalyst supported on certain silicas
US4180689 *Dec 11, 1978Dec 25, 1979The British Petroleum Company LimitedProcess for converting C3 -C12 hydrocarbons to aromatics over gallia-activated zeolite
US4350835 *Feb 19, 1981Sep 21, 1982Mobil Oil CorporationProcess for converting ethane to aromatics over gallium-activated zeolite
US4377504 *May 1, 1981Mar 22, 1983Phillips Petroleum CompanyCracking catalyst improvement with gallium compounds
US4415440 *Sep 8, 1982Nov 15, 1983Phillips Petroleum CompanyCracking catalyst improvement with gallium compounds
US4446013 *Nov 22, 1982May 1, 1984Phillips Petroleum CompanyCatalytic skeletal isomerization
US4566988 *Mar 21, 1984Jan 28, 1986Chevron Research CompanyProcess for the oxidation of gaseous hydrocarbons
US4891463 *Aug 5, 1988Jan 2, 1990Mobil Oil CorporationAromatization of aliphatics over a zeolite containing framework gallium
US4911823 *May 12, 1989Mar 27, 1990Mobil Oil CorporationCatalytic cracking of paraffinic feedstocks with zeolite beta
US4922051 *Mar 20, 1989May 1, 1990Mobil Oil Corp.Process for the conversion of C2 -C12 paraffinic hydrocarbons to petrochemical feedstocks
US4968650 *Apr 19, 1989Nov 6, 1990Mobil Oil CorporationZSM-5 catalysts having predominantly framework gallium, methods of their preparation, and use thereof
US4969987 *Nov 29, 1989Nov 13, 1990Mobil Oil CorporationIntegrated process for production of gasoline and ether
US5047070 *May 30, 1989Sep 10, 1991Mobil Oil CorporationIntegrated process for production of gasoline and ether from alcohol with feedstock extraction
US5160424 *Nov 13, 1990Nov 3, 1992Mobil Oil CorporationHydrocarbon cracking, dehydrogenation and etherification process
US8137535 *Jan 29, 2008Mar 20, 2012Kellogg Brown & Root LlcMethod for adjusting catalyst activity
US20050159633 *Dec 21, 2004Jul 21, 2005Ineos Chlor LimitedUse of alkali metal doped eta-alumina as methanol hydrochlorination catalyst
US20090192338 *Jul 30, 2009Pritham RamamurthyMethod for adjusting catalyst activity
EP0134328A1 *Aug 15, 1983Mar 20, 1985Mobil Oil CorporationTreatment of catalysts
EP0395345A1 *Apr 24, 1990Oct 31, 1990ARCO Chemical Technology, L.P.Production of olefins
U.S. Classification208/117, 585/407, 208/141, 502/355, 208/135, 585/415, 585/653
International ClassificationC10G11/04, B01J23/08, B01J27/10
Cooperative ClassificationC10G2400/30, B01J23/08, B01J27/10, C10G2400/20
European ClassificationB01J23/08, B01J27/10