Platinum containing catalysts are widely used in the oil refining and petrochemical industries, and are particularly important in a reforming process where paraffins, olefins and naphthenes are converted to aromatic compounds. Conventional reforming catalysts typically include one or more metals, most typically platinum, dispersed on a base, and may also include a binding agent for adding physical support to the base, and chloride to provide an acidic function. Typically, the catalyst base is alumina, but recently molecular sieve based catalysts have been found to be effective for reforming reactions.

Catalytic compositions containing zeolites are well known in the industry and recently the use of L-zeolites in combination with other specified catalytic components have been found to be particularly preferred for reforming. The aromatic compounds produced by such a catalytic conversion are valuable to a refiner due to their higher octane rating, and may be recovered from the reforming product for further processing and reaction in the petrochemical industries. The L-zeolite catalysts are particularly effective for converting C ₆

and Cη non-aromatic hydrocarbons which normally boil between 125°F and 225°F, to benzene and toluene.

In the reforming process, a hydrocarbon feedstock is passed through a catalyst-containing reactor in the presence of hydrogen at an elevated temperature. In the reactor and upon contact with the reduced or activated catalyst, some of the paraffins, olefins and naphthenes in the feedstock react to form a more desired, higher octane aromatic product. In the course of typical reforming operations, the catalysts will typically become deactivated due to the deposition of carbonaceous material or "coke" upon the catalyst, and/or scintering or poisoning of the catalytic metal particles.

Reforming catalysts that have been deactivated in this manner are typically regenerated by a method comprising a coke burning step, a platinum redispersion step, and a reduction step. In the coke burning step, the catalyst is contacted by an oxygen containing gas at elevated temperature to burn off coke deposits. In most cases, the maximum coke burn temperature exceeds 900°F. Platinum redispersion involves contacting the catalyst with a halogen compound and optionally oxygen and/or water at temperatures between 700°F and 1000°F. After platinum redispersion, the temperature is usually lowered and the reactor is purged with inert gas prior to starting catalyst reduction. The catalyst is then reduced by contacting with hydrogen.

Many variations of this method have been patented. The following patents apply specifically to the regeneration of reforming catalysts comprising a Group VII metal on a zeolite support: U.S. Patent No. 3,986,982 (Crowson et al) ; U.S. Patent No. 4,359,400 (Landolt et al) ; U.S.

Patent No. 4,493,901 (Bernard et al) ; U.S. Patent No. 4,810,683 (Cohn et al) ; U.S. Patent No. 4,914,068 (Cross et al) ; U.S. Patent No. 4,925,819 (Fung et al) ; U.S. Patent No. 5,106,798 (Fung et al) ; U.S. Patent No. 5,155,074 (Mohr et al) ; and U.S. Reissue 34,250 (Van Liersburg et al) . The methods in these patents have in common a coke burn at temperatures greater than or equal to 800°F and platinum redispersion using a halogen containing gas.

U.S. Patent No. 5,155,075 (Innes et al) describes a method for regeneration of a Pt-L-zeolite reforming catalyst wherein the coke burn is done at temperatures less than 780°F. The catalyst is then reduced with hydrogen while increasing the temperature to maximum between 900°F and 1000°F. Since the coke burn is done at low temperatures, platinum redispersion is not needed. The regeneration process is therefore halogen- free. U.S. Patent No. 5,073,529 also describes a halogen-free regeneration method, but does not limit the carbon burn temperature to less than 800°F.

U.S. Patent No. 5,270,272 (Galperin) describes a regeneration method wherein sulfur is removed from the catalyst by treatment with ammonia in nitrogen or hydrogen at very high temperatures. Since contacting the catalyst with ammonia at high temperatures causes scintering of the platinum, it is necessary to redisperse the platinum using halogen compounds prior to reduction.

Recently, it has been discovered that the fouling rate of certain zeolitic reforming catalysts can be greatly reduced by treating these catalysts in a reduced state at temperatures ranging from 1025°F to 1275°F. One such

high treatment procedure is described in U.S. Patent No. 5,382,353 (Mulaskey et al) . As a result of the high temperature treatment, the catalyst runs two to six times as long before regeneration is needed. Alternatively, the increased resistance to fouling makes it possible to operate at higher throughputs and/or lower hydrogen to hydrocarbon ratios. It also allows heavier feedstocks to be processed. Each of these process improvements has a substantial economic benefit.

Unfortunately, the regeneration processes of the prior art do not fully restore the fouling resistance of the high temperature treated catalyst. It has also been found that repeating the original high temperature treatment prior to the second reaction period usually leads to lower activity and a higher start of run temperature. The economic benefit of the high temperature treatment has thus far been mostly limited to the first cycle. A process which offers a successful and practical solution to the problem of regenerating such catalysts without substantial loss in activity or stability would be of great value to the reforming industry.

Accordingly, it is an object of the present invention to provide a process for regenerating a high temperature treated zeolite catalyst.

Another object of the present invention is to provide a process which can efficiently and effectively regenerate a high temperature treated zeolite without sacrificing catalytic activity or stability.

These and other objects of the present invention will become apparent upon a review of the following specification and the claims appended thereto.

SUMMARY OF THE INVENTION

In accordance with the foregoing objectives, the present invention provides a process for regenerating a high temperature treated reforming catalyst, which has been deactivated due to coke deposition, and which catalyst comprises at least one Group VIII metal supported on a zeolite base. For the purpose of this invention, a high temperature treated catalyst is defined as a catalyst that has been treated in an inert gas or reducing atmosphere at a temperature greater than or equal to 1025°F.

The process comprises the steps of

(a) contacting the catalyst with an oxygen containing gas at temperature and for a time sufficient to remove at least part of the coke deposited on the catalyst; then

(b) reducing the catalyst with a hydrogen containing gas; and

(c) treating the catalyst in an inert-gas or reducing atmosphere at temperatures in the range of 975°F to 1150°F.

It is preferred that the high treatment step be conducted in manner which limits the water concentration in the effluent gas to 200 ppmv or less.

The process of the present invention regenerates high temperature treated catalysts with a minimum loss of

activity or run length. Among other factors, the present invention recognizes that regeneration is different from activation. The temperature treatment is different in the regeneration than in the activation. Temperature treatments of 1150°F and above have been found to be extremely detrimental if the catalyst has been previously subjected to a treatment at temperatures of above 1025°F, and thus the treatment range for the final step of the present process is much lower than when treating fresh catalyst. The present invention permits one to run a reforming process for many cycles while utilizing a zeolite catalyst which was originally activated by a method comprising treatment with an inert gas or hydrogen containing gas in the temperature range of from 1025°F to 1275°F. The regeneration process of the present invention avoids any substantial lowering in the catalytic activity or stability of the high temperature treated catalyst. Such high temperature treated catalysts exhibit improved activity and a longer run life. The process of the present invention permits one to utilize these catalysts and take advantage of their improved activity and longer run life for numerous cycles.

BRIEF DESCRIPTION OF THE DRAWING

FIG. 1 of the Drawing is a plot of average catalyst temperature versus time representing the run of Example 6.

FIG. 2 of the Drawing is a plot of average catalyst temperature versus time representing the initial run of Example 7.

FIG. 3 of the Drawing is a plot of average catalyst temperature versus time representing the second run of Example 7.

FIG. 4 of the Drawing is a plot of average catalyst temperature versus time representing the run of Example 8.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS The catalyst regenerated in the process of the present invention has been deactivated due to coke deposition in the reforming process. The catalyst has also been previously subjected to a high temperature treatment at a temperature of at least 1025°F. Preferably, this previous high temperature treatment comprised treating the catalyst at a temperature in the range of from 1025°F to 1275°F, while maintaining the water level of the effluent gas below 200 ppmv.

More specifically, the catalyst is a large-pore zeolite charged with at least one Group VIII metal. The preferred Group VIII metal is platinum, which is more selective for dehydrocyclization and which is more stable under reforming reaction conditions than other Group VIII metals. The catalyst should contain between 0.1% and 5% platinum based on the weight of the catalyst, more preferably from 0.1% to 2.0%, and most preferably from about 1.0 to 1.5 wt %, e.g., about 1.2 wt %. The use of at least 1.0 wt % platinum is considered preferred as it helps the activity and stability of the catalyst in working with the naphtha feedstocks containing more than 5.0 wt % C ₉ ⁺ hydrocarbons.

The term "large-pore zeolite" is defined as a zeolite having an effective pore diameter of from 6 to 15 Angstroms. The preferred pore diameter is from 6.5 to 10 Angstroms. Type L zeolite, zeolite X, and zeolite Y, zeolite beta and synthetic zeolites with the mazzite structure are suitable large-pore zeolites for this operation. Zeolites with non-intersecting large pores such as zeolite-L and mazzite are thought to benefit most from the high temperature treatment. Type L zeolite is described in U.S. Patent No. 3,216,789.

Zeolite X is described in U.S. Patent No. 2,882,244. Zeolite beta is described in U.S. Patent No. 3,308,069. ZSM-4, described in U.S. Patent No. 4,021,447, is an example of a zeolite with the mazzite structure. Zeolite Y is described in U.S. Patent No. 3,130,007. U.S. Patent Nos. 3,216,789; 2,882,244; 3,130,007; 3,308,069; and 4,021,447 are hereby incorporated by reference to show zeolites useful in the present invention. The preferred zeolite is a type L zeolite.

Type L zeolites are synthesized largely in the potassium form. These potassium cations, however, are exchangeable, so that other type L zeolites can be obtained by ion exchanging the type L zeolite in appropriate solutions. It is difficult to exchange all of the original cations, since some of these cations are in sites which are difficult to reach. It may also be desirable at times to only partially exchange the potassium cations. The potassium may be ion exchanged, fully or partially, with an alkali or alkaline earth metal, such as sodium, cesium, lithium, rubidium, barium, strontium, or calcium. Preferably, in an exchange, the total amount of alkali or alkaline earth metal ions should be enough to satisfy the cation exchange sites of the zeolite or be slightly in excess.

It is preferred that the zeolite L contain exchangeable cations, at least 90% of which are selected from the group consisting of Li, Na, K, Rb, Cs, Ba and Sr ions or mixtures thereof.

An inorganic oxide can be used as a carrier to bind the large-pore zeolite. This carrier can be natural, synthetically produced, or a combination of the two. Preferred loadings of inorganic oxide are from 5% to 50% of the weight of the catalyst. Useful carriers include silica, alumina, aluminosilicates, and clays. The original high temperature treatment of the catalyst may occur at any time in the life of the catalyst. It is preferred that this treatment be carried out with fresh catalyst before use in the reforming process.

Preferably, the high temperature treatment used on the fresh catalyst occurs in the presence of a reducing gas such as hydrogen, as described in U.S. Patent No. 5,382,353, issued January 17, 1995, which is hereby expressly incorporated by reference in its entirety. Generally, the contacting occurs at a pressure of from 0 to 300 psig and a temperature of from 1025°F to 1275°F for from 1 hour to 120 hours, more preferably for at least 2 hours, and most preferably at least 4-48 hours. More preferably, the temperature is from 1050°F to 1250°F. In general, the length of time for the pretreatment will be somewhat dependent upon the final treatment temperature, with the higher the final temperature the shorter the treatment time that is needed.

In another embodiment, the catalyst can be treated using an inert gaseous environment in the temperature range of from 1025-1275°F, as described in copending U.S. Serial

No. (Attorney Docket No. 005950-442/T-

5123) , filed concurrently herewith, which is hereby expressly incorporated by reference in its entirety.

The preferred inert gas used is nitrogen, for reasons of availability and cost. Other inert gases, however, can be used, such as helium, argon and krypton, or mixtures thereof. The use of purely an inert gas atmosphere for the high temperature treatment allows one to avoid the problems inherent in using a reducing gas such as hydrogen.

The feed to the reforming process is typically a naphtha that contains primarily paraffins, olefins and naphthenes, generally having normal boiling points in the range of 100-400°F, and more preferably 160-350°F. This feed should be substantially free of sulfur, nitrogen, metals and other known poisons. These poisons can be removed by first using conventional hydrofining techniques, then using sorbents to remove the remaining sulfur compounds and water.

Because the catalyst of the present invention has been pretreated as previously described, it exhibits a longer run life with heavier feedstocks, e.g., containing at least 5 wt % C ₉ + hydrocarbons, than similar catalysts having been subjected to a different treatment. The catalyst obtained via the treatment of the present invention, therefore, makes it quite practical to process feedstocks containing at least 5 wt % C ₉ + hydrocarbons, and for example at least 10 wt % C _Q + hydrocarbons, with from 10-20 wt % C ₉ + hydrocarbons being preferred.

In the reforming process, the feed is contacted with the catalyst in either a fixed bed system, a moving bed system, a fluidized system, or a batch system. Either a fixed bed system or a moving bed system is preferred. In a fixed bed system, the preheated feed is passed into at least one reactor that contains a fixed bed of the catalyst. The flow of the feed can be either upward, downward, or radial. The pressure is from about 1 atmosphere to about 500 psig, with the preferred pressure being from abut 50 psig to about 200 psig. The preferred temperature is from about 800°F to about 1025°F. The liquid hourly space velocity (LHSV) is from about 0.1 hr ^"1 to about 10 hrs" ¹ , with a preferred LHSV of from about 0.3 hr" ¹ to about 5 hrs' ¹ . Enough hydrogen is used to insure a H ₂ /HC ratio of up to about 20:1. The preferred H ₂ /HC ratio is from about 1:1 to about 6:1. Reforming produces hydrogen. Thus, additional hydrogen is not needed except when the catalyst is reduced and when the feed is first introduced. Once reforming is underway, part of the hydrogen that is produced is recycled over the catalyst.

During the reforming reaction, the gradual accumulation of coke and other deactivating carbonaceous deposits on the catalyst will eventually reduce the activity of the catalyst and selectivity of the aromatization process. Typically, catalyst regeneration becomes desirable when from about 0.5 to about 3.0 weight percent or more of carbonaceous deposits are laid down upon the catalyst. At this point, it is typically necessary to take the hydrocarbon feed stream out of contact with the catalyst and purge the hydrocarbon conversion zone with a suitable gas stream. The catalyst regeneration method of the present invention is then performed either by

unloading the catalyst from the conversion zone and regenerating in a separate vessel or facility, or performing regeneration in-situ. Alternatively, the catalyst may be continuously withdrawn from the reactor for regeneration in a separate vessel, to be returned to the reactor as in a continuous catalytic reformer.

According to the catalyst regeneration process of the present invention, the initial step involves treating the catalyst with an oxygen containing gas to burn off coke deposits. It is preferred that the coke is burned from the catalyst by maintaining the catalyst at a temperature less than 800°F, and preferably less than 750°F. The burn step is preferably effected by contacting the deactivated catalyst with a gaseous mixture of oxygen and an inert gas. The oxygen is typically derived from air and an inert gas serves as a diluent, such that the oxygen concentration ranges from about 21 mole percent oxygen to a lower limit which for the practice of the present invention may be as low as 0.1 mole percent oxygen. The burn step is not limited to the use of air, however, and a higher level of oxygen may used in methods where oxygen is supplied in a more pure form such as from cylinders or other containing means. Typical inert gases useful in the low temperature coke burn step may include nitrogen, helium, carbon dioxide and like gases or any mixture thereof. Nitrogen is the preferred inert gas, however.

The regeneration gases should be substantially sulfur free as they enter they reactor, and preferably contain less than 100 parts per million by volume water. Because the oxygen content determines the rate of burn, it is desirable to keep the oxygen content low so as not to damage the catalyst by overheating and causing metal

agglomeration. It has been found desirable to keep the oxygen level in the inlet to the regeneration vessel between 0.2 to 4.0 mole percent during the coke burn step to avoid thermal damage to the catalyst, and still allow for the regeneration process to be accomplished in a reasonable amount of time.

Other conditions present during the coke burn step include a pressure sufficient to maintain the flow of the gaseous oxygen containing mixture through the catalyst zone such as a pressure of between 1.0-50.0 atmospheres and preferably from about 2 to about 15 atmospheres, and a gas hourly space velocity of about 100 to about 10,000 per hour, with a preferred value of about 500 to about 5,000 per hour.

It is also preferred that the burn step is conducted in a halogen free environment. By halogen free is meant that chlorine, fluorine, bromine or iodine, or their compounds including for example hydrogen chloride, are not added at any time during the catalyst regeneration process. In general, the coke burn step can follow along the lines of the low temperature regeneration process described in U.S. Patent No. 5,155,075, which is incorporated by reference expressly herein in its entirety.

Coke removal may also be done at temperatures above

800°F. In which case, a platinum redispension step using halogens will be required prior to reduction.

After the coke burnoff, the catalyst is reduced by contacting the catalyst with a reducing gas in the temperature range of from 300°F to 700°F. The temperature is then raised in a stepwise or ramping

fashion to complete reduction and drying. The reducing gas is preferably hydrogen, although other reducing gases can also be used. The hydrogen is generally mixed with an inert gas such as nitrogen, with the amount of hydrogen in the mixture generally ranging from 1 to 99% by volume. Preferred conditions for the initial reduction include a temperature in the range of about 400°F to about 600°F for a period of from about 0.1 to 10 hours. The pressure and gas rates utilized in the reduction step are preferably very similar to those described above with regard to the coke burn step.

Subsequent to reduction, the reduced catalyst is then treated at a temperature in the range of from about 975°F to less than 1150°F, and most preferably in the range of from 1000°F to about 1100°F. It is also most preferred that the water level of the effluent gas during the treatment of the catalyst at the high temperature is maintained below 200 ppmv, for otherwise the activity of the catalyst may be detrimentally effected.

It is also important that the temperature of this treatment not exceed 1150°F, and preferably 1100°F. For it has been discovered that when the temperature during this treatment reaches or exceeds 1150°F, catalyst activity is severely sacrificed. This high temperature treatment is not the same as the original high temperature treatment conducted on fresh catalyst, and therefore a different temperature profile must be observed and followed.

Treatment of the catalyst in the temperature range of from 975°F to less than 1150°F can be conducted in the presence of a reducing gas, such as hydrogen, or an

inert gas atmosphere. The preferred reducing gas is that of hydrogen, with the hydrogen generally being mixed with an inert gas, such as nitrogen. When an inert gas atmosphere is used, it is preferred that nitrogen be the inert gas, although other inert gases such as helium, argon or krypton, or mixtures of inert gases, can also be used.

The temperature for the final treatment is generally achieved by raising the temperature from the reducing step at a rate of between 5°F and 50°F per hour until the final treatment temperature is reached. More preferably, it is preferred that the temperature is increased at a rate of between 10°F and 25°F per hour. The temperature can be increased in a stepwise or ramping fashion. It is most preferred, particularly when the treatment range of 975°F is approached, that the temperature program and gas flow rates be selected to limit water vapor levels in the reactor to less than 200 ppmv, and preferably, less than 100 ppmv. This is particularly desirable when the catalyst bed temperature exceeds 1000°F.

During the final temperature treatment of the catalyst, it is preferred that the gas flow through the catalyst bed (GHSV) exceed 500 volumes per volume of catalyst per hour, where the gas volume is measured at standard conditions of one atmosphere and 60°F. From the standpoint of catalyst performance, the higher the gas velocity the better. GHSVs between 600 and 2000 h ^-1 are most preferred from a practical point of view.

To aid in maintaining the water level of the effluent gas below 200 ppmv in the final treatment of the catalyst, the inert or reducing gas entering the reactor

should contain less than 100 ppmv water. It is preferred that the gas contain less than 10 ppmv water. The effluent gas, may be passed through a drier containing a desiccant or sorbent such as 4 A molecular sieves. The dried gas can then be recycled to the reactor.

The length of the final treatment step can vary depending upon gas velocity, temperature, and catalyst particle size. Generally, however, the final treatment in the regeneration process will range from about 1 hour to about 120 hours, more preferably, for at least 2 hours, and most preferably in the range of from about 4 to 48 hours. In general, the length of time for the final treatment will be dependent upon the final treatment temperature. The higher the final temperature the shorter the time at final temperature needed. However, it is important that the temperature not reach 1150°F in the final treatment for otherwise catalyst activity and stability will be severely sacrificed. It is also important that the temperature be high enough and be maintained for a sufficient length of time to achieve an activity and stability approaching that of the original catalyst. Thus, the temperature range of from 1000°F to 1100°F is most preferred for the final treatment.

The invention will be illustrated in greater detail by the following specific examples. It is understood that these examples are given by way of illustration and are not meant to limit the disclosure or the claims to follow. All water measurements are in parts per million by volume (ppmv) .

EXAMPLE 1 A highly stable catalyst was obtained using the following activation procedure with fresh catalyst. A catalyst comprising 0.64% platinum on silica-bound, barium-exchanged, L-zeolite 1/16" diameter extrudates was charged to a pilot plant reactor. The catalyst was dried by circulating nitrogen at 70 psig and 1000 h" ¹ GHSV through the catalyst bed and a molecular sieve drier while heating the reactor to 500°F. Nitrogen circulation was continued until the water level in the reactor effluent dropped below 100 ppmv.

The catalyst was reduced by adding hydrogen to the recycle gas stream and increasing the total pressure to 100 psig. Thereafter, hydrogen was added to maintain pressure. The temperature was held at 500°F until the water concentration in the reactor effluent dropped below 100 ppmv. The catalyst then was heated at a rate of 10°F/h to 900°F. The temperature was held at 900°F until water in the reactor effluent fell again fell below 20 ppmv. The temperature was raised at a rate of 10°F/h to 1100°F. Above 900°F, the water concentration in the reactor was then less than 50 ppmv. The temperature was held for three hours at 1100°F before cooling to reaction temperature.

The catalyst thus activated was used under a variety of conditions for the conversion of naphtha to aromatics. By the end of the run, the catalyst activity loss corresponded to about a 70°F increase in the reaction temperature. When testing was completed, the catalyst was stripped of hydrocarbons with hydrogen and the reactor was purged with nitrogen and cooled to room temperature. The deactivated catalyst was ground into

20-40 mesh granules and thoroughly mixed. Portions of this catalyst were then regenerated in a microreactor to study the effects of carbon burn temperature and high temperature treatment temperature.

Each experiment was carried out as follows. A catalyst sample was heated to 500°F in nitrogen flowing at 50 psig and 8000 h" ¹ GHSV. A carbon burn was initiated by replacing the nitrogen with a 1.0 % oxygen in a nitrogen blend. The catalyst was heated 25°F/h to a final burn temperature between 700 and 800°F. The oxygen/nitrogen flow continued at this temperature for 20 hours to complete the carbon burn. The reactor was then cooled to 500°F.

The catalyst was then reduced with dry (< 10 ppmv water) hydrogen at 500°F, 50 psig, and 8000 h" ¹ GHSV. Hydrogen flow continued as the catalyst was heated 10°F/h to a final temperature between 900 and 1150°F. At temperatures above 975°F, the water concentration in reactor effluent was less than 100 ppmv. The catalyst was held at the final temperature for 20 hours and then cooled to reaction temperature.

Each regenerated catalyst sample was tested for the conversion of a light naphtha to benzene and toluene. Table 1 shows how the final carbon burn and hydrogen treatment temperatures affected catalyst performance. When the hydrogen treatment temperature was limited to 975°F or less, the catalyst deactivated at a much faster rate than when the hydrogen treatment was 1000°F or higher. The best results were obtained when the final carbon burn temperature was 700°F and the final hydrogen treatment temperature was 1000°F. A 750°F carbon burn was acceptable, but an 800°F burn temperature caused

significant catalyst deactivation. Surprisingly, the regenerated catalyst samples did not require as high a hydrogen treatment temperature as a fresh catalyst to produce a beneficial effect on catalyst stability.

Table 1. Effect of Regeneration Temperatures on Catalyst Performance

Final Final Wt % Yield Wt % Yield

Carbon Hydrogen Aromatics Aromatics

Burn Treatment after Two after Eight

Temp. , F Temp. , F Days Days

700 900 30 25

700 975 35 30

700 1000 36 36

750 1000 30 30

800 lOOO 24 20

700 1025 32 30

700 1050 31 30

700 1100 29 30

700 1150 29 30

Run Conditions: light naphtha feed, 10WHSV, 50 psig, 950°F, 5.0 hydrogen/naphtha feed molar ratio

EXAMPLE 2 An eighty cubic-centimeter portion of a catalyst comprising 0.64% platinum on silica-bound, barium- exchanged, 1/16-inch, L-zeolite extrudates was charged to a one-inch diameter tubular reactor. The catalyst was dried as in Example 1. The catalyst reduction and high temperature treatment were done with once-through hydrogen flowing at 2000 h ^"1 GHSV and 70 psig. The catalyst was reduced initially at 500°F, the catalyst temperature was raised lOF/h to 1100°F. The water concentration in the reactor effluent at temperatures above 900°F was less than 70 ppmv. Above 975°F, the

water concentration was less than 50 ppmv. The temperature was held at 1100°F for three hours before cooling to reaction temperature.

This catalyst was tested for the conversion of a light naphtha to benzene, toluene, ethylbenzene, and xylenes. The naphtha feed rate was 128 mL/h, the hydrogen to naphtha feed molar ratio was 3.0, and the reaction pressure was 100 psig. The reaction temperature was adjusted to maintain a 51.5 wt % aro atics concentration in the debutanized liquid product. The start-of-run average catalyst temperature for the target aromatics level was 847°F and the fouling rate was 0.010°F/h. By way of comparison, a catalyst activated with hydrogen at 500-900°F had a start of run temperature of 847°F and a fouling rate of 0.025°F. At the end of the run, the naphtha feed was stopped and the catalyst stripped of hydrocarbons with hydrogen. The catalyst was then purged with nitrogen and cooled to room temperature.

EXAMPLE 3 The catalyst from Example 2 was regenerated as follows. The catalyst was heated to 500°F as nitrogen was recirculated at 70 psig pressure through the reactor and recycle-gas drier. At 500°F, air was added to the recycle gas stream to initiate the carbon burn. The air feed rate was adjusted to maintain a 0.5% oxygen concentration at the reactor inlet. When oxygen appeared in the reactor effluent, the temperature was raised 25°F/h to 700°F. Upon reaching 700°F, the oxygen concentration was raised from 0.5 % to 1.0 %. After 24 hours the air feed was stopped. Nitrogen circulation continued while the reactor cooled to 500°F. At 500°F, the compressor was stopped and once-through nitrogen

flow started. The reactor pressure was 50 psig and the GHSV was 1000 h' ¹ .

Catalyst reduction was initiated at 500°F by slowly replacing nitrogen with hydrogen until the stream consisted entirely of hydrogen. The temperature was then increased from 500 to 1150°F at rate of lOF/h. The catalyst temperature was maintained at 1150°F for three hours and then allowed to cool to reaction temperature.

The regenerated catalyst was tested in the same way as the fresh high temperature treated catalyst in Example 2. Compared to the first cycle, the start-of- run temperature was 875°F versus 847°F and the fouling rate was 0.009°F/h versus 0.010°F/h. The 1150°F treatment caused the catalyst to loose a significant amount of activity (28°F) , but the catalyst still exhibited excellent stability.

EXAMPLE 4 The catalyst from Example 2 was regenerated a second time after completing the catalyst test in Example 3. This time the maximum temperature during the high temperature treatment step was 1000°F instead of 1150°F. After the second regeneration, the start of run temperature was 872°F compared to 875°F for the previous cycle and the fouling rate was unchanged. Thus, a 1000°F treatment maintained a low fouling rate without causing a further loss of start-of-run activity.

EXAMPLE 5 Eighty milliliters of a Pt-Ba-L-zeolite catalyst of the same type used in Examples one through four were charged to 1.0-inch diameter pilot plant reactor. The

catalyst was dried by heating to 500°F in flowing nitrogen. The reactor was at atmospheric pressure and the flow-rate was 3.0 ft ³ /h. The nitrogen flow continued at 500°F until the water concentration in the reactor effluent was less 100 ppmv. The catalyst was then reduced at 500°F by changing the gas to dry hydrogen and increasing the reactor pressure to 50 psig. The GHSV was adjusted to 4500 h" ¹ . After the water level in the reactor effluent again dropped below 100 ppmv, the temperature was increased from 500°F to 1100°F at a rate of 10°F/h. After holding for three hours at 1100°F, the reactor was cooled to reaction temperature.

The catalyst was tested for the conversion of a C ₆ -C ₇ naphtha feed to benzene and toluene. The start-of-run temperature was 852°F and the fouling rate was 0.003°F/h. This compares to 847°F and 0.020°F/h for the same catalyst and feed when the catalyst is activated with hydrogen at 500-900°F. When the run was completed, the catalyst was stripped of hydrocarbons with hydrogen. The reactor was then purged with nitrogen and cooled to ^' room temperature.

Later the catalyst was heated to 500°F in nitrogen flowing at rate of 3 SCF/min. At 500°F, air was added to the nitrogen at rate sufficient to give a reactor inlet oxygen concentration of 0.5%. After oxygen breakthrough, the temperature was raised 25°F/h to 700°F. The carbon burn was continued at 700°F with 1.0% oxygen for 24 hours. After the carbon burn, the catalyst was reduced at 500°F and heated to 1100°F in hydrogen following the same procedure used in the first cycle. The reforming process was then resumed under the original conditions.

The start-of-run temperature for the second cycle was 857°F and the fouling rate was 0.013°F/h. The 1100°F treatment resulted in only a 5°F loss in start-of-run activity compared to the first cycle. The fouling rate in the second cycle was better than achieved when the hydrogen treatment temperature is limited to 900°F, but was not as low as in the first cycle.

EXAMPLE 6 Eighty cubic centimeters of a catalyst comprising 1.2% platinum on K-Ba L-zeolite 1/16 inch extrudates were charged to a one-inch diameter reactor. The catalyst was dried by circulating nitrogen at 60 psig and 1000 hr-1 GHSV through the catalyst bed and a molecular sieve drier downstream of the catalyst bed. With the nitrogen gas recirculating, the reactor temperature was increased to 500°F at 25°F/hr. On reaching 500°F, nitrogen recirculation was continued until the concentration of water in the recirculating gas had decreased to 100 ppmv. The catalyst was then reduced by adding hydrogen to the recycle gas while maintaining 60 psig, and a recirculation rate of 1000 hr-1 GHSV. The reactor temperature was held at 500°F during the hydrogen addition until the recirculation gas had changed from 100% nitrogen to 95% + hydrogen. During the hydrogen addition, reactor pressure was held at 60 psig by excessing under pressure control some of the recycle gas (nitrogen+hydrogen) . In addition, to prevent against sudden increases in water concentrations in the recycle gas, hydrogen addition was stopped at any time that the concentration of water in the recycle gas increased to 400 ppmv. When the water concentration decreased to less than 400 ppmv, hydrogen addition was resumed. Once the nitrogen had been completely replaced with hydrogen and the water content of the recycle gas had decreased

to 100 ppmv, the reactor temperature was increased at lOF/hr to a temperature of 1100°F. On reaching 1100°F, this condition was held for 3 hours before cooling the reactor to reaction temperature.

A C ₆ -C ₉ naphtha containing 13.29% Ca ₉ ⁺ hydrocarbon was passed over the catalyst at 75 psig, 1.0 LHSV and a 5/1 H2/HC mole ratio. The specific components of the feed are described in Table 2 below.

TABLE 2

Feed Description

ASTM - D86, op

LV%, St 145

10 184

30 198

50 219

70 243

90 262

EP 295 gravity , API 65.8

Carbon No. distribution - wt%

C ₅ 1.82 c ₆ 27.72

C ₇ 22.77 c ₈ 33.77

C ₉ 13.29

^C 10 0.72

PNA - Wt %

P (n+i) 72.32 naphthenes 17.67 aromatics 9.37 unknown 0.64

Total 100.00

Feed was hydrofined to reduce sulfur content to acceptable levels.

The operating temperature was selected to achieve 83.5 wt% aromatics in the C ₅ ⁺ liquid. After the fresh catalyst had stabilized, the fouling rate was calculated to be 0.016°F/hr with an average start-of-run temperature of 864°F. The plot of average catalyst temperature versus time is shown in Figure 1.

After about 900 hours on stream, the naphtha feed was stopped and the catalyst was readied for regeneration.

EXAMPLE 7 The catalyst of Example 6 was regenerated by burning off the carbon deposited on the catalyst as described in Example 3 with the following exceptions. The reactor pressure was 85 psig. The reactor was heated up to 700°F at a 10°F/hr rate. On reaching 700°F and with the oxygen concentration at 0.5% to the reactor inlet, this condition was held for four hours. Following this step, the oxygen content was increased to 1%. After 24 hours at 700°F and 1% oxygen, the air feed was discontinued, the reactor was purged with nitrogen and cooled to 500°F.

The catalyst was then reduced in hydrogen as described in Example 6 except that the pressure was held at 60 psig and the catalyst was heated at 10°F/hr to a final temperature of 1050°F as compared to 1100°F in Example 6. Following the three hour hold at 1050°F, the catalyst temperature was reduced to reaction temperature.

The same C ₆ -C ₉ naphtha as used in Example 6 was passed over the regenerated catalyst at the same operating conditions as in Example 6. After the catalyst had stabilized, the fouling rate was 0.012°F/hr and the average start-of-run temperature was 867°F. The plot of the average catalyst temperature versus time is shown in Figure 2. The fouling rate of the regenerated high temperature hydrogen-treated catalyst was significantly lower than that obtained when the catalyst was fresh, namely 0.016°F/hr. In addition, the start-of-run temperature is only 3°F higher than with the fresh catalyst. In other words, the regenerated catalyst lost only 3°F of activity as a result of the regeneration and high temperature hydrogen treatment. Thus, surprisingly, the regenerated catalyst does not require as high a hydrogen treatment temperature as the fresh catalyst to equal or exceed the fresh catalyst stability.

At about 1000 hours on stream in the second cycle, the C ₆ -C ₉ naphtha feed was discontinued and replaced with a C ₆ -C ₇ naphtha. After 900 hours on stream with this new feed, the catalyst fouling rate was calculated to be 0.004°F/hr The start-of-run temperature with this feed was estimated to be about 845°F. The plot of the average catalyst temperature versus time is shown in Figure 3.

EXAMPLE 8 The catalyst of Example 7 was regenerated for a second time by burning off the carbon deposited on the catalyst. The regeneration was carried out as described in Examples 3 and 7. Following the removal of the deposited carbon, the catalyst was reduced in hydrogen and high temperature treated in hydrogen as described in

Example 7. In particular, as in Example 7, the final treatment temperature was 1050°F.

After cooling the reactor/catalyst to reaction temperature, the C ₆ -C ₇ naphtha used in Example 7 was passed over the catalyst. After 1000 hours on stream, the catalyst fouling rate was calculated to be 0.005°F/hr with a start-of-run average catalyst temperature of 855°F. The plot of average catalyst temperature versus time is shown in Figure 4. Thus, the second regeneration and the third high temperature hydrogen treatment resulted in a 10°F loss in initial catalyst activity relative to the second cycle and a slightly higher fouling rate than obtained in the second cycle, i.e., 0.005 versus 0.004°F/hr. Again, this shows that for the third cycle good catalyst performance was achieved without having to subject the catalyst to the same high hydrogen treatment temperature as the fresh catalyst.

While the invention has been described with preferred embodiments, it is to be understood that variations and modifications may be resorted to as will be apparent to those skilled in the art. Such variations and modifications are to be considered within the purview and the scope of the claims appended hereto.