DE69132272T2 - Regeneration process of a deactivated catalyst - Google Patents

Description

This invention relates to a process for activating a catalyst comprising at least one Group VIII noble metal component and a support, and also to a process for regenerating and reactivating coke-contaminated deactivated catalyst comprising one or more Group VIII noble metals and a support. These catalysts are frequently used in hydrocarbon conversion processes, particularly in catalytic reforming.

A number of materials have been used as hydrocarbon conversion catalysts in processes such as reforming, catalytic dewaxing, alkylation, oxidation and hydrocracking. Examples of catalysts useful in these processes include materials consisting of catalytically active Group VIII metal, typically platinum, and optionally rhenium or tin, supported on or impregnated into a support. Conveniently this support is alumina, although more recently zeolites have been used, for example X and Y type zeolites. L type zeolites have also been suggested as suitable supports, particularly cation exchanged L type zeolites.

Catalytic reforming in the presence of hydrogen is one of the most important hydrocarbon conversion processes. Catalytic reforming is a refinery process designed to increase the octane rating of naphtha. Typically, the naphtha feedstock is passed over a platinum-containing catalyst supported on a refractory material under reforming conditions, such as elevated temperatures and pressures, as is well known in the industry, in the presence of hydrogen gas at a hydrogen to hydrocarbon molar ratio of about 2 to 20. The reforming process involves several different types of reactions including isomerization, dehydrogenation of naphthenes to aromatics, dehydrogenation of paraffins to olefins, dehydrocyclization of paraffins and olefins to aromatics, hydrocracking of paraffins to gaseous hydrocarbons such as methane or ethane and inevitably the formation of coke, the latter depositing on the catalyst.

During the reforming process, the activity of the catalyst gradually decreases due to the build-up of coke until eventually regeneration and reactivation is required. Several stages are required for the regeneration and reactivation of the catalyst.

The first step is to regenerate the catalyst by removing the settled coke. Typically, the coke is removed by heating the catalyst in the presence of dilute oxygen to a flame front temperature of 430°C to 540°C. Some water vapor may be present in this regeneration step. US-A-4 354 925 discloses the use of a gaseous mixture of oxygen and carbon dioxide, the presence of this gas mixture enabling the use of a higher concentration of oxygen in the fuel gas, thereby enabling coke burn-off to be completed more rapidly. This regeneration step may be preceded by flushing the system with, for example, hydrogen or nitrogen gas to remove residual hydrocarbons.

The high temperatures used in the above regeneration result in agglomeration of the metal particles and, if the support is a zeolite, in removal of the metal particles from the zeolite channels and thus deactivation of the catalyst. Thus, after coke burn-off, the catalyst is usually subjected to a regeneration step in which the metal particles are redispersed (redispersed) on the support material. Typically this involves treating the catalyst with chlorine-containing gas or chlorine-releasing gas, usually in the presence of oxygen and water vapor, and is often referred to as an oxychlorination step.

Subsequently, the catalyst is treated with reducing gas, typically hydrogen, to reduce the metal chloride formed in the above oxychlorination step to elemental metal particles. The catalyst is then ready for reuse in the reforming process.

However, it has been found that, particularly with platinum-containing zeolite catalysts, the above regeneration and reactivation process often does not reactivate the catalyst to activation levels that correspond to or approach those exhibited by fresh catalyst.

However, it has been found that, particularly with platinum-containing zeolite catalysts, the above regeneration and reactivation process often does not reactivate the catalyst to activation levels that correspond to or approach those exhibited by fresh catalyst.

FR-A-2 642 330 and its counterpart GB-A-2 228 426 each disclose a process for regenerating catalyst (which may be reforming catalyst) containing metal, halogen and a support and used in a process carried out in at least two reactors in series through which the catalyst and the feed circulate in sequence, the first reactor through which the catalyst and the feed pass being at a pressure below 10 bar and the last reactor being at a pressure between 1 and 6 bar, the spent catalyst moving in a progressive manner downwards into a regeneration chamber in which it is successively applied to a first radial moving bed combustion zone, a second radial moving bed combustion zone, an axial moving bed chlorination or oxychlorination zone and a axial moving bed calcination zone, and

(a) the catalyst is treated in the first combustion zone at a pressure substantially equal to that in the first reactor at a temperature between 350 and 450ºC with a combustion gas based on inert gas containing 0.01 to 1 vol.% oxygen, the combustion gas coming from a washing zone as defined below,

(b) the catalyst in the second combustion zone at a pressure substantially equal to that in the first reactor and at a temperature at least 20°C higher than the temperature in the first combustion zone, in the presence of the gases from the first combustion zone and in the presence of of inert fresh gas containing up to 20 vol.% oxygen, so that the catalyst is in contact with gas containing 0.01 to 1 vol.% oxygen,

(c) the fuel gases are withdrawn from the second combustion zone and divided into two parts, and a first part containing 80 to 90 vol.% of the total combustion gases is mixed with the gases withdrawn from the oxychlorination zone defined below, the resulting mixture being passed to a washing loop and a drying zone, and a second part containing 20 to 10 vol.% of the total fuel gases is mixed with inert gas containing oxygen and with chlorine and/or one or more chlorinated compounds to form oxychlorination gas containing 1 to 4 vol.% oxygen,

(d) the catalyst in the oxychlorination zone is treated with the oxychlorination gas at a pressure of 3 to 8 bar and a temperature of between 350ºC and 550ºC and the gases are withdrawn from the oxychlorination zone to be mixed with the gases withdrawn from the second combustion zone as previously indicated,

(e) the catalyst is treated in the calcination zone at between 350 and 550°C at a pressure of 3 to 8 bar with at least a portion of the gases coming from the washing loop and a drying zone with a supplementary amount of fresh air and in such a way that the gases introduced into the calcination zone contain 1 to 10 vol.% oxygen and not more than 100 ppm water vapor, and

(f) the other part of the gases from the scrubbing loop in stage (a) is used as combustion gas.

WO-A-88/04574 discloses a process for redispersing platinum on a catalyst (e.g. reforming catalyst) comprising platinum, zeolite-L and inorganic oxide binder, in which

(a) the catalyst is oxychlorinated under oxychlorination conditions,

(b) the catalyst is contacted with inert gas, and

(c) the catalyst is contacted with dry hydrogen (preferably containing less than 100 ppm water).

In step (b), the inert gas is preferably nitrogen, preferably contains less than 1000 ppm water and the contact time is between 10 minutes and 90 minutes.

The present invention provides, in one aspect, a process for activating a reforming catalyst comprising at least one Group VIII noble metal component and a support, comprising the steps of

(a) CO₂ and CO are removed from the catalyst and

(b) the catalyst obtained in step (a) is chemically reduced in the absence of CO₂ and CO or under a partial pressure of CO₂ and CO not exceeding 0.025 kPa at a total pressure of 0.1 to 2 MPa.

The chemical reduction step (b) may comprise treating the catalyst with hydrogen-containing gas.

Step (a) may comprise purging with gas which is inert. Step (a) may comprise subjecting the catalyst to cyclic pressurization/depressurization with inert gas. The present invention, in another aspect, provides a process for regenerating and reactivating coke-contaminated reforming catalyst, the catalyst comprising at least one Group VIII noble metal component and a support, and the coke contamination results from use of the catalyst in a hydrocarbon reforming process, in which steps (c) coke is removed from the catalyst,

(d) the noble metal component(s) is/are redispersed on the support and (e) the catalyst is activated using the process described herein.

In step (d), the coke-depleted catalyst may be subjected to an oxychlorination step. Step (d) may be carried out in the presence of CO₂. The partial pressure of CO₂ in Level (d) may be at least 2 kPa, e.g. in the range 2 to 35 kPa.

The CO2 contacting of the catalyst may be terminated at least simultaneously with the termination of the contacting with halogen or a source thereof in the oxychlorination step (d).

In the process, the catalyst may be subjected to a stabilization step (f) after the oxychlorination step and before the activation step (e). Step (f) may comprise treating the catalyst with oxygen-containing gas. The catalyst may be contacted with CO₂ during the oxychlorination step (d). Contacting with CO₂ may be continued during the stabilization step (f).

The support may comprise a zeolite component. The zeolite component may comprise zeolite L.

Platinum may be a Group VIII noble metal component of the catalyst. The catalyst may comprise binders.

The invention also provides a catalytic hydrocarbon reforming process in which a hydrocarbon feedstock is contacted with a reforming catalyst under reforming conditions, in which the reforming catalyst is activated according to the process described herein or is regenerated and activated according to the process described herein.

In accordance with some embodiments, the invention provides a process for regenerating and reactivating coke-contaminated deactivated catalyst comprising at least one Group VIII noble metal, support and binder, which generally comprises

(c) substantially all of the coke is removed from the catalyst,

(d) the precious metal(s) are redispersed on the carrier,

(e') the catalyst is chemically reduced, wherein the redispersion step (d) is carried out by contacting the catalyst in a reaction vessel with a halogen or halide-containing gas and carbon dioxide gas, wherein the amount of carbon dioxide in the reaction vessel is substantially maintained at a partial pressure of at least 2 kPa during the period in which the catalyst is in contact with the halogen- or halide-containing gas.

We have found that the presence of carbon dioxide in the redispersion stage of the catalyst regeneration and reactivation process increases the activity and hence the performance of the catalyst. A catalyst which has undergone the process steps of the invention exhibits greater selectivity to form the desired aromatic compounds in a naphtha reforming process and also exhibits greater longevity, i.e. it retains sufficient activation levels over a longer period of time than catalysts regenerated by conventional techniques. The process of the invention offers particular advantages for the regeneration of Group VIII containing catalysts supported on zeolite material.

Detailed description of the invention

According to more specific embodiments of the present invention described below, the gaseous stream used in each stage has a residual percentage (which is not water, hydrogen, oxygen or a chlorine source) of an inert gas such as helium, argon or nitrogen that does not interfere with the process. Preferably, water is present in the gaseous stream from each stage.

Generally, any coke removal step (c) is carried out by contacting the deactivated catalyst with a gaseous stream containing 0 to 10 volume percent water, preferably 0.5 to 5 volume percent water, more preferably 1 to 3%, based on the total stream volume, and oxygen (usually in the form of an oxygen-containing gas such as air) at a temperature of about 380 to 540°C. Preferably, this first step is carried out at two temperatures, the first being a lower temperature range of about 380 to 480°C and the second being a higher temperature range of about 480 to 520°C. The O₂ treatment at the lower temperature is preferably for a longer period than the second O₂ treatment. The exact duration of the heating depends on the temperature used, but is generally in the range of 1 to 10 hours or more. The amount of oxygen used is generally 0.1 to 25% by volume, preferably 0.2 to 15% by volume, especially 0.5 to 12% of the gas stream.

Any redispersion step (d) is carried out using halogen or halide-containing gas (hereinafter generally referred to as "halogen gas") and carbon dioxide. The source of the halogen gas may be, for example, chlorine gas, carbon tetrachloride or hydrogen chloride. The halogen gas is usually mixed with oxygen and preferably also with water vapor and this redispersion step is usually referred to as an oxychlorination step. Examples of suitable oxychlorination steps are:

(i) heating the catalyst to a temperature of about 400 to about 530°C for up to 10 hours in the presence of a gaseous stream comprising zero to 10 volume percent water and a chlorine source in the presence of oxygen,

(ii) heating the catalyst to a temperature of about 400 to about 530°C for up to 10 hours in the presence of a gaseous stream comprising zero to 10 volume % water and a chlorine source, in the presence of hydrogen (except if HCl is the chlorine source, then the hydrogen is optional) and then heating the catalyst to a temperature of about 400 to 540°C for up to 7 hours in the presence of a gaseous stream comprising 0 to 10 volume % oxygen and essentially no water, or up to 5 hours in the presence of a gaseous stream comprising 0 to 10 volume % oxygen and greater than zero to 10 volume % water,

(iii) combining steps (ii) and (i) above in that order without carrying out the subsequent heating in step (ii),

(iv) heating the catalyst to a temperature of about 400 to about 530°C for up to 10 hours in the presence of a gaseous stream comprising zero to 10 vol.% water and an effective amount of chlorine.

The source of carbon dioxide may, for example, be pure carbon dioxide gas fed directly to the reactor in which the catalyst is regenerated, or a mixture of carbon dioxide and inert gases. The carbon dioxide source may also be mixed with the other gases used in the oxychlorination stage, typically halogen gas, oxygen and optionally water vapor. Alternatively, the source of carbon dioxide may be a carbon dioxide-releasing compound, which may be gaseous, liquid or solid, although gaseous is preferred for ease of introduction into the reactor.

To obtain improved catalytic performance, the carbon dioxide must be maintained at a partial pressure of at least 2 kPa in the reactor substantially during the period that the catalyst is in contact with the halogen gas. Conveniently, the introduction of carbon dioxide gas is started at about the same time as the introduction of the halogen gas and is terminated at about the same time as the termination of the addition of the halogen gas. The carbon dioxide can be fed into the reactor, for example, as a continuous gas stream or as gas pulses, provided that the minimum concentration is maintained. Preferably, a continuous stream is used. The carbon dioxide gas exiting the reactor can be recycled and fed back into the reactor, thereby providing an economical reactivation process.

Preferably, the partial pressure of carbon dioxide in the reactor during the oxychlorination step (i.e. when halogen gas is present) is from 2 kPa to 35 kPa, especially from 4 kPa to 15 kPa. It is preferred to introduce carbon dioxide gas into the reactor as a mixture of carbon dioxide gas with inert gases, for example nitrogen or helium.

Although the present invention is not intended to be limited to any particular theory, it is believed that the presence of carbon dioxide increases the extent of reactions between the ha logen gas and any binder present in the catalyst, such as alumina, and thereby inhibits the formation of metal halides, such as platinum chloride, on the binder material. This therefore leads to a better distribution of the platinum particles within the channels of the zeolite L in platinum-zeolite L catalysts.

After completion of the oxychlorination step, it is generally preferred to remove any excess halogen and to stabilize the metal particles on the support surfaces. Typically, this is achieved by continuing to contact the catalyst with oxygen after the halogen gas flow has ceased, for example by heating the catalyst to a temperature for up to 7 hours in the presence of a gaseous stream comprising 0 to 10 volume percent oxygen and essentially no water, or for up to 5 hours in the presence of a gaseous stream comprising 0 to 10 volume percent oxygen and greater than zero to 10 volume percent water. This additional step is not required when the redispersion step (ii) above is carried out because it is already integrated into this step. If desired, the carbon dioxide flow can be continued during the oxygen stabilization step, after which the flow is stopped.

After oxychlorination, the catalyst must be chemically reduced to convert the noble metal halide formed during oxychlorination to the elemental metal. Generally, this reduction is accomplished by contacting the catalyst with a reducing gas such as hydrogen at elevated temperatures, optionally in the presence of water vapor. For example, the catalyst is heated to a temperature of about 350 to about 530°C for up to 10 hours in the presence of a gaseous stream comprising zero to 10 volume percent water and a hydrogen source. The hydrogen concentration is generally from 0.5 to 100 volume percent, preferably from 2 to 20 percent.

We have also found that the presence of carbon dioxide or carbon monoxide during the reduction step can adversely affect the performance of the resulting catalyst.

Therefore, the reduction is preferably carried out in the presence of a minimal amount, if any, of carbon dioxide and carbon monoxide. The amount of carbon dioxide and carbon monoxide gases present during the hydrogen reduction step should be not more than 0.025 kPa partial pressure. It is particularly preferred to remove carbon dioxide and carbon monoxide prior to the reduction step when the catalyst support is a zeolite L. If the system is purged with oxygen after the oxychlorination step as described above, this should effect the removal of the carbonaceous gases. However, it is also preferred to remove any remaining oxygen prior to introducing the hydrogen gas. However, whether the oxygen treatment has been carried out or not, it is preferred to purge the system with an inert gas, for example nitrogen or helium, prior to commencing the reduction step. An alternative way to remove the carbon monoxide or dioxide prior to the reduction stage is by cyclic pressurization/depressurization with an inert gas such as nitrogen that is free of carbon monoxide or dioxide. By pressurizing the reactor with the inert gas, the concentrations of carbon monoxide and dioxide are reduced by dilution. This process can be coupled with gas purging as described above.

In all of the stages of catalyst regeneration and reactivation according to the invention, the pressure of the reactor is 0.1 to 2 MPa, conveniently atmospheric pressure. The gas flow rates are generally in the range of about 1 to about 300 ml per gram of catalyst per minute.

The catalyst support may be, for example, an inorganic oxide such as alumina, titania, zinc oxide, magnesium oxide, thoria, chromia or zirconia or the like, a zeolite, for example faujasite, mordenite, zeolite X, Y or L, a clay such as china clay, kaolin, bentonite, diatomaceous earth or other silicon-based materials such as silica gel or silicon carbide or a mixture of one or more of the above. Preferably, the support is alumina or a zeolite, particularly a zeolite and especially zeolite type L. Zeolites of type L can be defined as synthetic zeolites which crystallize in the hexagonal system. They have channel-shaped pores whose diameter varies from about 7 to 13.4 μm and they can occur in the form of cylindrical crystals with an average diameter of at least 0.5 μm and an aspect ratio of at least 0.5 (as described, for example, in US-A-4,544,539, the disclosure of which is incorporated herein by reference) as well as in other shapes and sizes. Zeolite L typically has the general formula:

0.9 to 1.3 M2/nO:Al&sub2;O&sub3;:XSiO&sub2; : yH2O

in which M represents an exchangeable cation, n is the valence of M, y is any value from zero to about 9, and x is about 5.2 to about 6.9.

A more complete description of zeolite L is given in US-A-3,216,789.

Type L zeolites are conventionally prepared so that M in the above formula is potassium. See, for example, US-A-3,216,789 and US-A-3,867,512. The potassium can be ion exchanged, as is well known, by treating the zeolite in an aqueous solution containing other cations. However, it is difficult to exchange more than 75% of the original potassium cations because some cations occupy sites in the zeolite structure which are almost inaccessible. At least 75% of the exchangeable cations are selected from lithium, sodium, potassium, rubidium, cesium, calcium and barium. More preferably, the cation is sodium, potassium, rubidium or cesium, especially potassium, rubidium or cesium, most preferably potassium. Optionally, the exchangeable cations may consist of mixtures of the above-mentioned Group IA cations or mixtures of Group IA cation and barium or calcium cations. These mixtures of cations can be obtained, for example, by treating the zeolite L with an aqueous solution containing, for example, B. a rubidium and/or cesium salt, and then washed to remove excess ions. This ion exchange treatment can be repeated to effect further ion exchange, albeit to a lesser extent.

The zeolite catalysts used in the processes of the present invention can be prepared by loading the zeolite with one or more Group VIII metals. Preferably, the zeolite is processed into shaped particles, e.g., pellets, extrudates, spheres, granules, or the like, with an inert binder such as alumina, silica, or clays prior to loading. Alumina, particularly γ-alumina, is a preferred binder. References to the diameter of the catalyst particles in the present specification should be construed as references to the diameter of the shaped particles, which is preferably from about 1/32" (0.79 mm) to about 1/4" (6.35 mm). Preferably, the metal is or includes platinum, typically about 0.3 to about 1.5 weight percent Pt based on the weight of the zeolite, and is loaded onto the zeolite by a method as described in U.S. Patent No. 4,568,656.

In the loading process of US-A-4,568,656, the zeolite, typically as a pellet or matrix with the binder, is contacted with an aqueous loading solution containing a platinum salt and a non-platinum metal salt to give a pH of between 8 and 12.5, aged to disperse the platinum salt in the pores of the zeolite, dried and then calcined.

The zeolite L catalysts can be prepared at different pH levels and we have found that the addition of carbon dioxide during the redispersion stage of catalyst regeneration and reactivation is particularly beneficial when platinum deposition on the zeolite L has been carried out at relatively low pH (around pH 8 to 9) rather than at higher pH values, although somewhat improved catalyst performance is also demonstrated when the zeolite catalysts are prepared at these higher values.

The Group VII noble metals necessary for catalytic activity are those metals from Group VIII of the Periodic Table of the Elements selected from osmium, ruthenium, rhodium, iridium, palladium and platinum. Preferably the metals used here are platinum, rhodium or iridium and most preferably platinum and examples of suitable bimetallic combinations are platinum-iridium and platinum-palladium. The metals may be present in any desired combination. Rhenium, a Group VIIB metal, may also be present as long as at least one Group VIII noble metal is present, an example being platinum-rhenium.

The amount of Group VIII noble metal present in the catalyst is an effective amount and depends, for example, on the required catalyst activity, the ease of uniform dispersion and the type of catalyst support. For zeolites, the crystal size limits the effective catalyst loading since highly loaded zeolite crystals having a large dimension parallel to the channels can easily lead to pore clogging during operation as the noble metal agglomerates inside the channels. In general, however, the amount of metal present is in the range of about 0.1 to 6 wt.%, preferably 0.1 to 3.5 wt.%, more preferably 0.1 to 2.5 wt.% and most preferably 0.2 to 1 wt.% of the catalyst. Furthermore, for zeolites, the amount of metal present is generally about 0.1 to 2.0 weight percent of the catalyst when the average size of the zeolite crystallites parallel to the channels is greater than about 0.2 microns, and about 1.0 to 6.0 weight percent when the average zeolite crystallite size parallel to the channels is not greater than about 0.2 microns.

The Group VIII noble metals can be introduced by, for example, ion exchange, impregnation, carbonyl decomposition, adsorption from the gas phase, introduction during zeolite synthesis and adsorption of metal vapor. The preferred technique is ion exchange. In some cases, e.g. when the metal(s) is/are introduced by an ion exchange process, it is preferred to reduce the residual acidity of the support material by treating the catalyst, which has previously been reduced with hydrogen, with an aqueous solution of an alkali base such as potassium carbonate. This treatment neutralizes any hydrogen ions formed during the reduction of the Group VIII noble metal ions by hydrogen.

The activation process of the present invention can also be applied to fresh catalysts to improve the activity of the fresh catalyst in naphtha reforming processes. By "fresh catalyst" is meant a catalyst that has never been used in a naphtha reforming process and is thus not contaminated with coke. Thus, in another aspect, the present invention provides a process for activating a fresh catalyst comprising at least one Group VIII noble metal, support and binder, which comprises

(a) the catalyst is treated in a reaction vessel with a halogen- or halide-containing gas in the presence of carbon dioxide, and

(b) the catalyst is chemically reduced.

The preferred conditions and embodiments of this other aspect of the invention are as described above for the process for reactivating coke-contaminated catalyst.

It has also been found that the reduction of the catalyst prior to its use in a reforming process, whether or not it has undergone activation or reactivation in the presence of carbon dioxide, is advantageously carried out in the presence of between zero and 0.025 kPa partial pressure of carbon dioxide and carbon monoxide. In particular, the reduction is carried out in the substantial absence of these carbonaceous gases. If a greater concentration of these gases is present, the activity of the catalyst during reforming is substantially reduced. This applies to the reduction of both fresh and regenerated catalysts.

Once the catalyst has been reduced, it is advantageous not to expose it to carbon dioxide and carbon monoxide in a amount greater than 0.025 kPa partial pressure before the catalyst is used in a reforming process.

The invention will now be illustrated by the following examples. Unless otherwise stated, all parts and percentages are by weight for solids and liquids and by volume for gases. Gas compositions have been described with the major components, the balance of the composition comprising an inert gas such as nitrogen or helium.

example 1

A catalyst was prepared as follows:

A zeolite L having a composition, expressed in moles of pure oxide, of 0.99 K₂O:Al₂O₃:6.3 SiO₂:x H₂O, cylindrical in shape and an average particle size of about 2 to 2.5 µm was prepared according to the technique described in Example 1 of US-A-4,544,539. Thus, an alkaline synthesis gel was prepared by dissolving 23.4 g of aluminum hydroxide by boiling in an aqueous solution of 51.2 g of potassium hydroxide pellets (86% pure KOH) in 100.2 g of water to form solution A. After dissolution, any water loss was corrected. A separate solution, Solution B, was prepared by diluting 225 g of colloidal silica (Ludox HS 40) with 195 g of water. Solutions A and B were mixed for two minutes to form a gel and just before the gel became completely rigid, 224 g of the same was placed in a Teflon-lined autoclave which was preheated to 150ºC and kept at that temperature for 72 hours to induce crystallization. The solid zeolite was then separated.

The zeolite was then formed into 1/16" (1.55 mm) extrudates with γ-alumina as a binder. 0.64 wt% platinum was then loaded into the zeolite extrudates using Pt(NH3)Cl2 salt, with the amount of KCl and KOH calculated according to the method disclosed in U.S. Patent No. 4,568,656 to give a final loading pH of 10.0.

The Pt/KL zeolite was then calcined at 350°C in a reactor at atmospheric pressure with a gas stream of 10% dry O2 at a flow rate of 500 ml/min for 2 h, after which 3% water was added to the gas stream and the treatment continued for an additional 2 h. The catalyst was then reduced with a gas stream of 10% 82 and 3% H2O at a flow rate of 500 ml/min for 3 h at a temperature of 350°C. The resulting fresh catalyst is referred to as Catalyst A.

The deactivation produced by coke burn-off in the first stage of catalyst regeneration was simulated by treating a 2 g sample of catalyst A with 10% O2 and 3% H2O at a flow rate of 500 ml/min in a reactor at atmospheric pressure for 2 h at 510°C. The deactivated catalyst was then oxychlorinated using conventional conditions. Thus, the catalyst was treated with 0.35% O2, 10% O2 and 3% H2O at a flow rate of 500 ml/min for 5 h at 510°C. The HCl flow was stopped and the catalyst stabilized by treatment with 10% O2 and 3% water for 1 h. The catalyst was then reduced with 10% H₂ and 3% H₂O at a flow rate of 500 ml/min for 2 h at 510°C. The resulting comparative catalyst is referred to as Catalyst B.

The deactivation and reactivation of another sample of 2 g of Catalyst A was repeated as for Catalyst B, except that CO2 was introduced into the reactor during the deactivation step and the CO2 flow was continued during the oxychlorination step and the subsequent oxygen stabilization step and stopped thereafter. The CO2 was fed into the reactor with the other gases and the CO2 concentration was maintained at 4% (4 kPa partial pressure). In addition, the oxychlorination time was 2.5 h and the subsequent stabilization and reduction were carried out at 350°C. The CO2 was purged from the reactor using an inert gas purge prior to the reduction step. The resulting catalyst is referred to as Catalyst C.

The performance of catalysts A, B and C was each evaluated by measuring the benzene yield produced from a reforming process with 3-methylpentane. Reforming was carried out at a H₂/3-methylpentane ratio of 6, a space velocity of 20 w/w/hr, a temperature of 510°C and a pressure of 7.24 bar gauge (105 psig). The results are given in Table 1. Table 1

The results show that when the catalyst is regenerated under normal conditions, as in Catalyst B, the activity of the catalyst in terms of aromatics production drops below that of fresh catalyst (Catalyst A), especially after a number of hours have elapsed. However, when the catalyst is treated with CO2 during the oxychlorination step, as in Catalyst C, the activity of the catalyst is significantly improved, in this case even exceeding that of the fresh untreated catalyst.

Example 2

A sample of Pt/KL zeolite catalyst was taken from a reforming reactor. The catalyst comprised 0.64% platinum on zeolite KL extrudates, with the platinum loaded onto the zeolite at a pH of 11 to 12. The catalyst had been in operation for more than 1000 h (i.e., used in the reforming process to convert naphtha) and had been regenerated three times in situ using conventional regeneration methods. This used catalyst is referred to as Catalyst D.

A 2 g sample of Catalyst D was regenerated in the same manner as described for Catalyst B in Example 1, except that the oxychlorination time was 2.5 hours. The resulting comparative catalyst is referred to as Catalyst E. Another 2 g sample of Catalyst D was regenerated in the same manner as described for Catalyst C in Example 1, except that the CO2 flow was stopped when the HCl flow was stopped and the oxygen stabilization and hydrogen reduction steps were carried out in the absence of CO2 at 510°C. The resulting catalyst is referred to as Catalyst F.

The performance of Catalyst D, E, F and a sample of the fresh commercial catalyst, designated Catalyst D¹, was evaluated using a 3-methylpentane reforming procedure as described in Example 1. The results are given in Table 2. Table 2

The results show that the used catalyst D is significantly deactivated compared to the fresh catalyst D1. The catalyst regenerated under normal conditions, comparative catalyst E, is reactivated to some extent, but a significant improvement is obtained when the catalyst has been reactivated in the presence of CO2.

Example 3

A Pt/KL zeolite catalyst was synthesized as described in Example 1, except that the platinum was added at a pH of 8.75 was loaded onto the zeolite extrudates. This catalyst is referred to as Catalyst G.

A sample of Catalyst G was deactivated and then reactivated in the same manner as Catalyst B in Example 1, except that the gas stream in the deactivation step also contained 10% CO2 (10 kPa partial pressure). The CO2 was purged from the reactor prior to the oxychlorination step. This catalyst is referred to as Catalyst H.

Another sample of Catalyst G was deactivated and reactivated in the same manner as Catalyst F in Example 2. This catalyst is referred to as Catalyst I.

The performance of catalysts G, H and I was evaluated using a 3-methylpentane reforming process as described in Example 1 and the results are given in Table 3. Table 3

The results show a very significant improvement in catalytic performance when CO2 is present in the oxychlorination stage, but not when CO2 is present only during deactivation (corresponding to the coke burning stage).

Example 4

A Pt/KL zeolite catalyst was synthesized as described in Example 1, except that the platinum was loaded onto the zeolite extrudates at pH 10.3. This catalyst is referred to as Catalyst J.

2 g of catalyst J was deactivated and then reactivated in the same manner as catalyst B in Example 1. This catalyst is referred to as catalyst K.

Another sample of catalyst J was deactivated and then reactivated in the same manner as catalyst F in Example 2. This catalyst is referred to as catalyst L. The performance of catalysts J, K and L was evaluated using a 3-methylpentane reforming procedure as described in Example 1 and the results are given in Table 4. Table 4

Again, the results show a very significant improvement in catalytic performance when CO₂ is present in the oxychlorination step. In addition, transmission electron spectroscopy (TEM) showed that both catalysts K and L contained highly dispersed Pt clusters of size 6 to 10 Å. However, there was a clear difference between these two catalysts. Platinum was detected on the Al₂O₃ binder of catalyst K, but was absent on the Al₂O₃ binder of catalyst L. This suggests that the presence of CO₂ in the oxychlorination step, the interaction between the Al₂O₃ binder and platinum chloride species is reduced and as a result the redistribution of Pt on the Al₂O₃ binder during oxychlorination is minimized.

Example 5

A sample of the fresh commercial catalyst (Catalyst D¹) of Example 2 was deactivated and then reactivated in the same manner as Catalyst F in Example 2, i.e., with inclusion of CO₂ in the deactivation and oxychlorination steps. This catalyst is referred to as Catalyst M.

Another sample of catalyst D¹ was deactivated and then reactivated as for catalyst M, except that 10% CO₂ (10 kPa partial pressure) was also included in the reduction step. This catalyst is referred to as catalyst N.

Another sample of catalyst D1 was deactivated and reactivated as for catalyst B in Example 2, i.e. under normal conditions except that 10% CO2 (10 kPa partial pressure) was present in the reduction step. This catalyst is referred to as catalyst O.

Another sample of catalyst D¹ was deactivated and reactivated as for catalyst O, except that the stabilization and reduction steps were each carried out at 350ºC. This catalyst is referred to as catalyst P.

Finally, catalyst D¹ was reduced with 10% H₂, 3% H₂O and 4% CO₂ (4 kPa partial pressure) at a temperature of 410°C for 2 h. This catalyst is referred to as catalyst Q. The performance of the above catalysts was evaluated using a 3-methylpentane reforming procedure as described in Example 1 and the results are given in Table 5. Table 5

The results show a significant loss of catalyst performance when CO2 is present in the reduction stage and the reduction temperature is between 350 and 510ºC, independently of the conditions of previous regeneration stages. TEM studies of catalysts N and O show the presence of platinum agglomerates of size 10 to 30 11.

Example 6

3 g of the fresh catalyst D1 from Example 2 were heated to 510°C in 10 vol% O2 in helium at a flow rate of 500 ml/min. After the catalyst reached 510°C, CO2 was added to the gas stream to give 1 vol% CO2 (1 kPa partial pressure) and after 5 minutes the oxygen was turned off. The remaining oxygen was purged from the reactor with the 1 vol% CO2 in helium at 500 ml/min for 10 minutes. Hydrogen was then added to give 10 vol% H2, keeping the total flow at 500 ml/min. This treatment was continued for 1.5 h. This catalyst is referred to as catalyst R.

An additional 2 g of catalyst D1 from Example 2 was heated to 400°C in 10 vol% O₂ in helium at a flow rate of 500 ml/min. After the catalyst reached 400°C, CO₂ and water were added to the gas stream to give 0.2 vol% CO₂ (0.2 kPa partial pressure) and 3 vol% water, respectively. The temperature was then raised to 510°C. The catalyst was treated under the above gas mixture for 35 minutes before the oxygen was turned off. The remaining oxygen was purged from the system for 10 minutes before hydrogen was added to give 10 vol% hydrogen. The reduction was carried out for 1.5 hours. This catalyst is referred to as Catalyst 5.

Another 3 g of catalyst D1 was first reduced at 510°C with 10 vol% hydrogen in helium at a flow rate of 500 ml/min for 1 h. Then CO2 was added to give 0.05 vol% (0.05 kPa partial pressure) and this treatment was continued for 1.5 h. This catalyst is referred to as catalyst T.

An additional 2 g of catalyst D1 from Example 2 was treated by the same procedure as catalyst T, except that in the H2 and CO2 stage, CO2 was present at 0.025 vol% (0.025 kPa partial pressure) and the gas flow rate was 1000 ml/min. This catalyst is referred to as catalyst U.

The performance of catalysts R, S, T and U was evaluated using a 3-methylpentane reforming procedure as described in Example 1 and the results are given in Table 6. Table 6

The results again show that CO2 is detrimental to the catalyst when present in the reduction stage of the Pt/KL zeolite catalysts. TEM studies of catalyst S showed that platinum particles with cluster sizes of 10 to 25 k were agglomerated. There is still a slight loss of activity even at 0.025 vol% (0.025 kPa partial pressure), although this level of loss is more acceptable. The above results also show that exposure of pre-reduced catalyst to CO2-containing hydrogen still leads to catalyst deactivation (catalysts T and U). Therefore, after catalyst regeneration, CO2 should be purged to a low level prior to hydrogen reduction. The exclusion of CO2 from the treatment gas is also important in the activation of fresh catalyst.

Example 7

2 g of fresh catalyst D¹ from Example 2 were heated to 200ºC in helium at a flow rate of 500 ml/min. The catalyst was dried under these conditions for 1 h. The temperature of the catalyst was then raised to 350ºC. Hydrogen was added to give 10 vol.%. The catalyst was reduced for 1.5 h. Then the gas mixture was CO was added to the mixture to give 0.5 vol.% CO (0.5 kPa partial pressure). This treatment was continued for 1.5 h. CO was turned off and was stripped from the catalyst with 10 vol.% hydrogen at 350ºC for 1.5 h. This catalyst is referred to as Catalyst V.

Another 2 g of catalyst D1 from Example 2 was treated in the same manner as catalyst V except that the CO concentration was 0.1 vol% (0.1 kPa partial pressure) and the gas flow rate in the H2 and CO stages was 1000 ml/min. This catalyst is referred to as catalyst W. Another 2 g of catalyst D1 from Example 2 was heated in 10 vol% O2 in helium at a flow rate of 500 ml/min to 400°C. After 1 h at 400°C, the temperature of the catalyst was increased to 510°C. Oxygen was purged for 10 min before CO was added to give 0.1 vol% (0.2 kPa partial pressure). Hydrogen was then added to give 10 vol%. This treatment was continued for 1.5 h. This catalyst is called catalyst X.

The performance of catalysts V, W and X was evaluated using a 3-methylpentane reforming procedure as described in Example 1 and the results are given in Table 7. Table 7

The results show that CO present in the hydrogen reduction stage deactivates the catalysts in the temperature range 350ºC to 510ºC. A catalyst will suffer more deactivation if the H₂ and CO treatment is carried out at a higher temperature (compare catalyst X and W). A reduced catalyst will also be deactivated when exposed to H₂ and CO. deactivated (catalysts V and W). TEM studies of catalyst X show that platinum particles were agglomerated, and particle sizes up to 40 Å were observed. Therefore, when reducing a fresh catalyst, CO should be purged to low concentration and the gas stream must be CO-free.

Example 8

3 g of catalyst D1 from Example 2 was heated to 200°C in helium at a flow rate of 500 ml/min. The catalyst was dried under these conditions for 1 h. Oxygen was added to give 10 vol% and the temperature of the catalyst was raised to 510°C. Then the oxygen was turned off and the reactor was purged with helium for 10 min. CO2 was added to the gas stream to give 1 vol% CO2 (1 kPa partial pressure). This treatment was continued for 1.5 h. This catalyst is referred to as catalyst Y. 2 g of catalyst D1 from Example 2 was heated to 350°C in 10 vol% O2 in helium at a flow rate of 500 ml/min. The catalyst was treated under these conditions for 20 minutes. Then the oxygen was turned off and the reactor was purged with helium for 30 minutes. CO2 was added to the gas stream to give 1 vol% (1 kPa partial pressure). This treatment was continued for 1.5 h. This catalyst is referred to as Catalyst Z.

2 g of catalyst D1 from Example 2 were heated to 200°C in helium at a flow rate of 500 ml/min. The catalyst was dried under these conditions for 1 h. Hydrogen was added to the gas stream to give 10 vol% H2. Reduction was carried out for 1.5 h. Hydrogen was then turned off and purged from the reactor with helium for 1 h. CO was added to the gas stream to give 0.5 vol% CO (0.5 kPa partial pressure). This treatment was continued for 1.5 h. This catalyst is referred to as catalyst AA.

2 g of catalyst D¹ in Example 2 were treated as catalyst AA, except that the CO concentration was 0.1 vol.% (0.1 kPa partial pressure) at a gas flow rate of 1000 ml/min. This catalyst was designated catalyst BB.

The performance of catalyst Z, AA and BB was evaluated using a 3-methylpentane reforming process as in Example 1 and the results are given in Table 8. Table 8

The results show that at low temperature (350ºC) the exposure of the catalyst to CO₂ in the absence of hydrogen does not lead to catalyst deactivation (catalyst Z). However, when the exposure temperature is increased to 510ºC, significant catalyst deactivation occurs (catalyst Y). In contrast, the exposure of the reduced catalyst to CO at low temperature (350ºC) in the absence of hydrogen still leads to catalyst deactivation.

BeisDiel 9

4 g of reforming catalyst obtained from Ketjen containing 0.6 wt% Pt in 1/16" (1.59 mm) Al2O3 extrudates (Catalyst CC) was heated to 510°C under 20 vol% O2 in helium at a flow rate of 500 ml/min. The catalyst was treated under these conditions for 1 h. Oxygen was turned off and the reactor was purged with helium for 10 min. CO and H2 were added to the gas stream to give 4 vol% CO (4 kPa partial pressure) and 20 vol% H2. This treatment continued for 2 h. This catalyst is referred to as Catalyst DD.

Fresh catalyst CC and "used" catalyst DD were evaluated using hydrogen chemisorption to measure the dispersion of the Pt particles. These two catalysts were first reduced at 475°C in 100 vol% hydrogen at a flow rate of 1000 ml/min for 2 h. They were then evacuated at 450°C for 1 h. Hydrogen chemisorption was carried out at 35°C. The detailed procedure and apparatus used for hydrogen chemisorption can be obtained from J. Am. Chem. Soc., 100, 170 (1978), by SJ Tauster, SF Fung and R. L Garten. This technique uses a measurement of the volume of hydrogen irreversibly adsorbed by one gram of catalyst. From this an atomic ratio of hydrogen atoms to metal atoms, H/Pt, can be calculated. The results are shown in Table 9. Table 9

The results show that exposure of the catalyst to CO in the reduction step also deactivates Pt/Al₂O₃ catalysts. The lower H/Pt ratio suggests that fewer Pt atoms are surface atoms and that the Pt clusters have increased in size.

Claims (18)

1. A process for activating a reforming catalyst comprising at least one Group VIII noble metal component and a support, in which in steps (a) CO₂ and CO are removed from the catalyst and (b) the catalyst obtained in step (a) is chemically reduced in the absence of CO₂ and CO or under a partial pressure of CO₂ and CO not exceeding 0.025 kPa at a total pressure of 0.1 to 2 MPa. 2. A process according to claim 1, wherein the chemical reduction step (b) comprises treating the catalyst with hydrogen-containing gas. 3. A process according to claim 1 or claim 2, wherein step (a) comprises purging with gas which is inert. 4. A process according to any one of claims 1 to 3, wherein in step (a) the catalyst is subjected to cyclic pressurization/depressurization with inert gas. 5. A process for regenerating and reactivating coke-contaminated reforming catalyst, wherein the catalyst comprises at least one Group VIII noble metal component and a support, and the coke contamination results from use of the catalyst in a hydrocarbon reforming process comprising the steps of (c) removing coke from the catalyst, (d) redispersing the noble metal component(s) on the support, and (e) activating the catalyst using a process according to any one of claims 1 to 4. 6. Process according to claim 5, wherein in step (d) the coke-depleted catalyst is subjected to an oxychlorination step. 7. A process according to claim 5 or claim 6, wherein step (d) is carried out in the presence of CO₂. 8. A process according to claim 7, wherein the partial pressure of CO₂ in step (d) is at least 2 kPa. 9. A process according to claim 7 or claim 8, wherein the partial pressure of CO₂ in step (d) is in the range 2 to 35 kPa. 10. A process according to any one of claims 7 to 9, wherein the CO₂- contacting of the catalyst is terminated at least simultaneously with the termination of the contacting with halogen or a source thereof in the oxychlorination step (d). 11. Process according to one of claims 6 to 10, in which the catalyst is subjected to a stabilization step (f) after the oxychlorination step and before the activation step (e). 12. A process according to claim 11, wherein in step (f) the catalyst is treated with oxygen-containing gas. 13. A process according to claim 11 or claim 12, wherein the catalyst is contacted with CO₂ during the oxychlorination step (d) and the contacting with CO₂ is continued during the stabilization step (f). 14. A method or process according to any one of claims 1 to 13, wherein the support comprises a zeolite component. 15. A method or process according to claim 14, wherein the zeolite component comprises zeolite L. 16. A method or process according to any one of claims 1 to 15, in which platinum is a Group VIII noble metal component of the catalyst. 17. A method or process according to any one of claims 1 to 16, wherein the catalyst comprises a binder. 18. A catalytic hydrocarbon reforming process in which a hydrocarbon feedstock is contacted with a reforming catalyst under reforming conditions, wherein the reforming catalyst is either (a) is activated according to the method according to one of claims 1 to 4 and/or the method according to one of claims 14 to 17, or (b) is regenerated and activated according to the process according to any one of claims 5 to 17.