PROCESS FOR OBTAINING ALKYL-NAPHTHENICS

Abstract

The present invention addresses to a process for the production of alkyl-naphthenics for use as diesel and/or aviation kerosene (JET A-1), whose process involves the alkylation of olefins with monoaromatics and subsequent hydrogenation to alkyl-naphthenics. The process and catalysts of the present invention allow the regeneration of the acidic catalyst with a hydrogenating function and full recovery of its activity with hydrogen hot stripping. The catalyst is used for the formation of intermediate alkyl-aromatics and can also be used in the subsequent hydrogenation to alkyl-naphthenics.

Claims

1. A process for obtaining alkyl-naphthenics, comprising: (a) alkylating aromatics with olefins using an acid-supported alkylation catalyst with at least one metal from Group 10 of the Periodic Table of Elements, plus at least one metal from Group 9, and/or Group 7; (b) regenerating the activity of the alkylation catalyst by contacting the alkylation catalyst with H.sub.2 at a temperature higher than the alkylation reaction performed in step (a); (c) repeating step (a) at least once using the regenerated alkylation catalyst produced in step (b).

2. The process according to claim 1, wherein the acid support of the alkylation catalyst is selected from at least one of aluminosilicates, amorphous or crystallines.

3. The process according to claim 2, wherein the acid support of the catalyst is selected from silica-aluminas, large pore zeolites in acidic form, such as ferrierites, chabazites, Y, US-Y, RE-Y, ZSM-5, ZSM-12, NU-86, mordenites, ZSM-22, NU-10, ZBM-30, ZSM-11, ZSM-47, ZSM-35, IZM-2, ITQ-6, IM-5, SAPO (silico-alumino-phosphates), Beta zeolite, MCM-22, MCM-56, phosphated or silanized molecular sieves treated with siloxanes, clays, pillared clays, mixed metallic oxides, acidic ion exchange resins, sulfonated silicas, and phosphated niobium.

4. The process according to claim 1, wherein the contents of metals from Group 10 are from 0.1 to 5% w/w and metals from Group 9 and/or Group 7 are from 0.05 to 2% w/w.

5. The process according to claim 4, wherein the contents of metals from Group 10 are from 0.2 to 1% w/w and metals from Group 9 and/or Group 7 are from 0.1 to 0.5% w/w.

6. The process according to claim 5, wherein the contents of metals from Group 10 are 0.6% w/w and metals from Group 9 and/or Group 7 are 0.2% w/w.

7. The process according to claim 1, wherein the effluent from the alkylation step is sent to the hydrogenation reaction of alkyl-aromatics to alkyl-naphthenics, using a hydrogenation catalyst.

8. The process according to claim 1, wherein the operating conditions of the alkylation reaction comprise a temperature between 100 and 400° C.; a pressure between 30 and 200 bar (3 and 20 MPa); and an LHSV between 0.1 and 10 h.sup.−1.

9. The process according to claim 8, wherein the operating conditions of the alkylation reaction comprise a temperature between 150 and 350° C.; a pressure between 30 and 100 bar (3 and 10 MPa)); and an LHSV between 0.5 and 4 h.sup.−1.

10. The process according to claim 9, wherein the operating conditions of the alkylation reaction comprise a temperature between 200 and 300° C.; a pressure of 60 bar (6 MPa); and an LHSV between 1 and 2 h.sup.−1.

11. The process according to claim 1, wherein the catalyst regeneration reaction is performed with super-atmospheric hydrogen, preferably with the same operating pressure as the alkylation.

12. The process according to claim 1, wherein the alkylation catalyst is contacted with H.sub.2 for between 4 h and 48 h.

13. The process according to claim 12, wherein the catalyst contact with H.sub.2 is between for 8 and 36 h.

14. The process according to claim 13, wherein the catalyst contact with H.sub.2 is for 24 h.

15. The process according to claim 1, wherein the regeneration step is performed using H.sub.2 is with a temperature between 250 and 500° C.

16. The process according to claim 15, wherein the regeneration step is performed using H.sub.2 is with a temperature between 350 and 450° C.

17. The process according to claim 16, wherein the regeneration step is performed using H.sub.2 with a temperature of is up to 550° C.

18. The process according to claim 1, wherein the regeneration of the alkylation catalyst performed in step (b) is carried out in the presence of the alkylation product with an excess of hydrogen, resulting in alkyl-naphthenics.

19. The process according to claim 1, wherein a) the at least one metal from Group 9 comprises rhodium; and/or b) the at least one metal from Group 7 comprises rhenium.

Description

BRIEF DESCRIPTION OF DRAWINGS

[0055] The present invention will be described in more detail below, with reference to the attached figures which, in a schematic way and not limiting the inventive scope, represent examples of the embodiment thereof. In the drawings, there are:

[0056] FIG. 1 illustrating the process scheme of the invention, with at least 3 reactors, where there are represented the olefinic charge (1), the aromatic charge (2), the alkylation reactor in operation (10), the alkylation product (11), the fractionator (20), the recycle of unreacted aromatics (22), the non-olefinic lights (21), the alkylate (23), the hydrogen for hydrogenation (3), the alkyl-aromatics hydrogenation reactor (30) and the alkyl-naphthenic product (31), as well as the regenerating hydrogen (4) and the regenerating reactor (40), loaded with the alkylation catalyst claimed in the present invention;

[0057] FIG. 2 illustrating the results of tests of different catalysts for alkylation, comparing the conversion of olefins with the yield in aromatics;

[0058] FIG. 3 illustrating the deactivation comparison for the various alkylation catalysts of the invention;

[0059] FIG. 4 illustrating the deactivation for catalyst A, with recoveries of activity by sequences of hydrogen hot stripping and oxide regeneration;

[0060] FIG. 5 illustrating the deactivation for catalyst C, with recoveries of activity by hydrogen hot stripping sequences.

DETAILED DESCRIPTION OF THE INVENTION

[0061] The process of producing aviation kerosene (JET A-1) and diesel from charges containing olefins and aromatics according to the present invention and illustrated in FIG. 1 consists of: [0062] (a) carrying out the alkylation reaction of aromatics with olefins with the acid-supported alkylation catalyst with at least one metal from Group 10 of the Periodic Table of Elements, plus at least one metal from Group 9, such as the Rh, and/or Group 7, like Re; [0063] (b) carrying out the reaction of regenerating the activity of the alkylation catalyst by contacting H.sub.2 at a temperature higher than the alkylation of step (a); [0064] (c) reusing the alkylation catalyst for step (a).

[0065] Olefins and aromatics can be present in different charges, fed to the same process, such as LPG from the FCC unit, and the reformate from the Catalytic Reform Unit (CRU). Alternatively, some charges may contain olefins and aromatics already in their composition, such as NFCC.

[0066] Useful charges existing in the refinery, which can be used in the process of the present invention, are charges containing olefins and/or aromatics in any proportion. FCC gas (containing olefins such as C2= and C3=), LPG from FCC (C4=), NFCC, LCO, coke naphtha (NLK), coke diesel, thermally cracked or pyrolysis gasolines/naphthas can be mentioned.

[0067] The conversion of olefinic and aromatic streams in the gasoline range makes room for the unreacted portions of these streams, of paraffinic and naphthenic nature, to be reprocessed at the FCC, for example, to produce more olefins and aromatics. This combination of FCC alkylation, for example, can essentially zero gasoline production in a refinery, offering maximum flexibility for the refining plant, significantly increasing diesel and JET A-1 production.

[0068] The present invention is characterized by the acid site reaction of a heterogeneous catalyst of aromatics with olefins (and dienes, when present).

[0069] Charge treatments and purifications can be used, such as providing adsorption means for nitrogenous, oxygenated, polar compounds in general, present in the charges. Another problematic component that leads to blockage of catalyst pores are dienes, which polymerize easily, even more so in the presence of oxygen.

[0070] However, given the regeneration capacity facilitated by the present invention, it is possible to operate without treatment of charges, simply by adjusting the frequency of regeneration operations.

[0071] Alternatively, a step of selective hydrogenation of dienes can be envisaged. In a possible configuration of the invention, the catalyst of the present invention itself, under liquid phase conditions and lower temperature, can hydrogenate the charge dienes to olefins, prior to contacting the same catalyst in a higher temperature alkylation condition. Catalysts for selective hydrogenation are known in the art, usually supported metals from Group 10 of the Periodic Table.

[0072] Different modes of operation of alkylation, regeneration, and hydrogenation are possible, to obtain the highest yield in the alkylation and hydrogenation reactions, accommodating the deactivation of the acidic function of the catalyst.

[0073] It is possible to alkylate the charges to a still acceptable level of catalyst activity, stop the reactor charge and align H.sub.2, increasing the alkylation temperature to the desired regeneration temperature, and realign the charge. The hydrogenation can further be carried out on the same catalyst, and the activity is recovered with the hydrogenation itself. It is further possible to carry out a regeneration step in a shorter time just to recover the metallic sites of hydrogenation capacity, before hydrogenation, and allow the oligomers to be removed during the hydrogenation step. At the end of the hydrogenation, the catalyst can be operated again for alkylation, and an additional step of hot stripping with H.sub.2 may occur before aligning the alkylation charge.

[0074] Alternatively, the hydrogenation catalyst of the alkyl-aromatics may be different from the alkylation catalysts.

[0075] The degree of hydrogenation depends on the destination of the product. As diesel, the total hydrogenation of the monoaromatics present is preferable, mainly to improve the density and cetane. As JET A-1, the product may not be hydrogenated, or only partially, depending on the amount of stream that will compose the JET A-1 pool.

[0076] It is possible to operate the process with only one reactor, discontinuously.

[0077] It is more interesting, however, to provide means of having more than one reactor or operating bed. Thus, in a preferred arrangement of the present invention, while one reactor regenerates, the other alkylates, and a third hydrogenates. Or with 2 reactors and a larger LHSV in the hydrogenation, it is possible to regenerate a bed, and use the deactivated catalyst for the alkylation for the hydrogenation and resume the alkylation in the regenerated catalyst.

[0078] The preferred operation is with 3 reactors, while a reactor alkylates, a second reactor regenerates and a third reactor hydrogenates. There may also be a fourth reactor carrying out selective hydrogenation of dienes from the olefinic charge, prior to alkylation.

[0079] In the case of 3 reactors, as the regeneration times are shorter than the alkylation times, it is possible to use the regenerated catalyst for other functions, to work in line or in parallel, to use as a guard bed, with less conversion, or another function, such as cracking non-aromatic products to generate more olefins or a hydro-alkylation step. In hydro-alkylation, with aromatic charge plus sub-stoichiometric hydrogen generates olefinic alkyl-naphthenics, which can be alkylated to the remaining aromatics, generating diaryl-alkyl-aromatics.

[0080] Preferably, when there is a mostly olefinic charge separated from the aromatic charge, it is possible to inject the olefins along the beds of a reactor, with the reactor having at least 2 beds. The more beds, the smaller the deactivation, but with an increase in the complexity of the unit, there is an optimum to be determined by economic considerations.

[0081] The reactor may or may not have product recycle. Preferably the reactor has product recycle. Product recycle reduces the need for cooling between reactor beds, since the reaction is exothermic. Furthermore, it reduces the concentration of olefins and the undesirable side reactions of formation of oligomers from light olefins in the gasoline range. Another advantage of recycling is that it increases the amount of aromatics with more than one alkyl (dialkyl-aromatics, trialkyl-aromatics), increasing the boiling point, quality and quantity of the product. An additional advantage of promoting more than one alkylation of the same aromatic is being able to convert a greater amount of olefinic charge. Typically, the reactor recycle can be represented by the ratio between the amount of product that is fed back to the reactor inlet divided by the reactor charge. It can be from zero to 20, preferably from 0.1 to 2. Furthermore, before recycling there can be a separator. With a separator, which can be a flash or a set of flashes or a distillation or adsorption unit, only the unreacted aromatics are fed back to the reactor.

[0082] A higher content of aromatics in relation to the olefinic charge is advantageous, reducing the formation of oligomers when the olefinic charge is light (e.g.: C4=), disfavoring the formation of oligomers with higher yield in the gasoline range, being selective in the formation of alkylates in the range of JET A-1 and diesel.

[0083] In the hydrogenation step, product recycle is also advantageous, not only because of the decrease in exothermicity. It may be interesting to dimension the recycle in order to allow the reactor operation in liquid or supercritical phase in the hydrogenation step, provided by the liquid recycle, without the need for a gas recycle compressor, sending only the H.sub.2 of chemical consumption to the unit. This allows higher reaction kinetics and higher rates of mass transfer in the reactor and less complexity of the unit, the pumping of liquid being more easily implemented than that of gas, as is known in the art.

[0084] Typical alkylation temperatures are temperatures from 100 to 400° C., preferably 150 to 350° C., more preferably 200 to 300° C., in general. Some catalysts, however, can operate at higher temperatures. What limits the temperature, however, to less than 500° C., is the possibility of sintering the catalyst metals. It is possible and desirable to start with high conversion and increase the temperature over the run time to extend the campaign time before regeneration.

[0085] Typical temperatures for hydrogenation of alkyl-aromatics in catalysts of a metal from Group 10 are 200 to 400° C., preferably 200 to 300° C. Above 300° C., the chemical equilibrium of hydrogenation is already evident, when increases in temperature mean less hydrogenation, under typical conditions of operating pressure, usually less than 100 bar (10 MPa). In addition, higher pressures of up to 200 bar (20 MPa) can be used.

[0086] Desirable pressure conditions are charge dependent. It is preferable to maintain the alkylation pressure above the critical pressure of the mixture. In the case of mixing toluene with LPG, the desirable pressure is greater than 55 bar (5.5 MPa). The critical temperature is around 250° C.; so, in most of the operation, the deposition of oligomers in the reactor will be reduced due to the higher diffusivities in supercritical medium. In practice, pressures greater than 30 bar (3.0 MPa) are preferable, preferably in the range of 60 bar (6 MPa), and pressures of up to 100 bar (10 MPa) are sufficient. For hydrogenation, it was found that in the present invention maintaining the same operating pressures as the alkylation allowed for the desired hydrogenation of alkyl-aromatics to alkyl-naphthenics.

[0087] The LHSV (volume of charges fed per reactor volume per hour) depends on the nature of the charges, pressure conditions, temperature and desirable campaign time before a regeneration step. A typical LHSV is from 0.1 to 10 h.sup.−1, preferably from 0.5 to 4 h.sup.−1, more preferably from 1 to h.sup.−1, for both hydrogenation and alkylation, although typically the LHSV conditions of the hydrogenation may be greater than those of alkylation.

[0088] The typical operating times before regeneration is required are at least 2 days to 1 month, typically 4 days to 2 weeks. Too long before regeneration can build up polymers on the surface in a way that makes it difficult to access the metal sites needed for catalyst regeneration. Also, too long time between regenerations can mean too low LHSV, and larger reactor sizes for a given charge, which is undesirable.

[0089] The typical regeneration conditions are higher than those employed in alkylation, typically from 250 to 500° C., more preferably from 350 to 450° C., and not higher than 550° C. The regeneration pressure can be the same or less than that used in alkylation. Greater pressures prove unnecessary. More preferably, the reactor in the regeneration step operates in a down-flow mode, in order to facilitate the flow of the liquid that previously wet the catalyst. The amount of required hydrogen is small, being 1 volume of H.sub.2 under normal conditions of temperature and pressure, per reactor per minute, preferably 10 volumes of H.sub.2 per reactor volume per minute, which is equivalent to a GHSV of at least 60 at 600 h.sup.−1, which may be higher depending on the need to heat the catalyst bed under the conditions necessary for the regeneration of the catalyst of the present invention.

[0090] The catalyst of the present invention contains both acidic and hydrogenating functions. The catalyst has a hydrogenating function and has a metal from Group 10 of the Periodic Table, preferably Pt and/or Pd, plus at least one metal from Group 9, such as Rh, and/or Group 7, such as Re.

[0091] The contents of metals from Group 10 are typically 0.1 to 5% w/w, more preferably 0.2 to 1% w/w, most preferably 0.6% w/w. Higher metal contents are unnecessary for complete catalyst regeneration, and decrease the availability of acidic sites.

[0092] The contents of metals from Group 9 and/or Group 7 are typically 0.05 to 2% w/w, more preferably 0.1 to 0.5% w/w, more preferably 0.2% w/w.

[0093] Preferably, the metals are prepared with precursors without chlorine or any other halides in the composition, which will add chlorine content to the catalyst.

[0094] Several catalysts of acidic nature can be used in the present invention, such as alumino-silicates, amorphous or crystallines. In general, silica-aluminas, large-pore zeolites in acidic form, such as ferrierites, chabazites, Y, US-Y, RE-Y, ZSM-5, ZSM-12, NU-86, mordenites, ZSM-22, NU-10, ZBM-30, ZSM-11, ZSM-47, ZSM-35, IZM-2, ITQ-6, IM-5, SAPO (silico-aluminum-phosphates), Beta zeolite, MCM-22, MCM-56, molecular sieves can also be phosphated or silanized (treated with siloxanes), clays, pillared clays, mixed metallic oxides, acidic ion exchange resins, sulfonated silicas, phosphated niobium.

[0095] FIG. 1 presents a preferred arrangement of the present invention. A stream containing olefins (1) is sent to a reactor (10), part of which is mixed with the charge (2) containing aromatics. In reactor (10), there are at least 2 beds, with part of the olefin injection being carried out after the first bed. Further, in addition to charges (1) and (2), a recycle (22) of product aromatics, obtained from the fractionation of the product stream (11) in a fractionator (20), is added. The bottom product (23) of the fractionator follows to an aromatic hydrogenation unit (30), where hydrogen (3) is added for the reaction, obtaining alkyl-naphthenics (31), or a mixture of alkyl-naphthenics and alkyl-aromatics, in case of partial hydrogenation. While reactors (10) and (30) are dedicated to alkylation and hydrogenation, a reactor (40) is regenerated by the hydrogen stream (4).

[0096] Whereas the scheme of FIG. 1 is preferred, other operating schemes are possible, for example alkylation, product accumulation and hydrogenation in the same reactor, followed previously or later by a step of regeneration of the acidic function of the catalyst.

[0097] Preferably, the streams sent to the hydrogenation step contain little sulfur, preferably below 500 ppm, in order to allow the use of metals from Group 10 for the aromatics hydrogenation step. Otherwise, the hydrodesulfurization reaction (HDS) in catalysts such as sulfided CoMo and NiMo is known to those skilled in the art. While sulfur removal is unfavorable in the case of olefinic streams due to undesired saturation of olefins, it can be used after alkylation, before the hydrogenation step, once the olefins are converted. The same catalyst could be used for HDS and hydrogenation, but separation is preferable in a subsequent step of hydrogenation of aromatics after removal of sulfurs, since the activity of sulfided catalysts for hydrogenation of mono-aromatics is low. The alkylation step is not significantly affected by the presence of sulfur, and sulfur-containing charges can be processed. The presence of other contaminants, however, such as nitrogenous ones, can decrease the time of the alkylation campaign, and it can be advantageous to previously remove at least part of these compounds by means known in the art, such as adsorption, washing of the stream, etc. In one scheme of the present invention, as the regeneration is fast compared to the time of the alkylation campaign, the regenerated bed can be used as a trap, at temperatures lower than the alkylation, and be regenerated again before the alkylation step itself.

[0098] In addition to the fixed bed schemes described in the present invention, fluidized beds or transported beds can be used. However, such a reaction scheme is unnecessary, since the recovery nature of the alkylation activity of the present invention allows for simpler fixed bed operation, and the greater difficulty of providing means for solids movement is unnecessary.

[0099] In the particular condition of the invention of increasing the temperature from the alkylation condition to the regeneration condition, it can be done by processing aromatic or paraffinic charge, up to the desired temperature, or by heating the hydrogen stream itself, or even the mixture from the hydrogen stream with inert stream, such as paraffinic C4. Means for heating, achieving and maintaining the regeneration condition are known in the art, and various schemes can be employed without departing from the regeneration claim of the present invention. As with heating, lowering the temperature is also employed by means known in the art in order to process the alkylating charge after regeneration.

EXAMPLES

[0100] The following examples are presented in order to illustrate some particular embodiments of the present invention, and should not be interpreted as limiting the same. Other interpretations of the nature and mechanism of obtaining the components claimed in the present invention do not change its novelty.

[0101] Experiments were carried out using refinery charges and model compost. To facilitate the analysis of the products, as aromatics, toluene was used, and as olefins, liquefied petroleum gas (LPG), containing mostly C4, with a total of 59.9% w/w in olefins, coming from the FCC unit.

[0102] The tests were conducted in an automated benchtop unit (PID). LPG and toluene were mixed in line, with independent pumps. The unit had N.sub.2 flow pressurization at the top of the separator vessel. Thus, the unit was pressurized upstream without contacting the catalyst with the gas, being able to maintain the desired pressure from the beginning of the tests. Also, before entering the gas-liquid separator, the reactor effluent, after cooling to room temperature, went through a loop to the chromatograph, for in-line analysis, without loss of light. A chromatograph with a mass detector and FID was used to identify and quantify the products. Gaseous effluent was also analyzed and quantified, and no significant amounts of light were formed in addition to those already present in the charge.

[0103] 5 ml of catalyst were used for each run, diluted in 5 ml of carborundum (SiC.sub.2). The catalysts when extruded were broken in length one by one, maintaining the diffusion size (length not lesser than the diameter), for reasons of hydrodynamics and mass transfer. The packaging, particle and reactor diameter ratios and minimum bed length for a high conversion (in the range of 95%) followed scale-down criteria for trickle-bed and liquid phase reactors, to guarantee representativeness of the larger scale even in a micro reactor.

[0104] Typical alkylation temperatures were used, from 90 to 360° C. The pressures used aimed to maintain the liquid phase and, preferably, in a condition close to critical or supercritical. Critical point estimation using process simulator for a typical charge composition (50 vol % LPG and 50 vol % toluene) showed that the critical pressure was approximately 55 bar (5.5 MPa) and the critical temperature above 250° C. LHSV conditions were varied from 0.5 to 4 h.sup.−1.

[0105] In all tests the catalytic bed was initially fed with aromatic (toluene) before feeding the specified flow rate of olefin.

Example 1: Test of State-of-the-Art Supports, without Addition of Hydrogenating Function

[0106] Representative catalysts of 5 classes of acidic catalysts were tested. A commercially available divinyl-benzene macroporous resin (DVB resin) in the H form (various sources such as Duolite C20, Duolite C26, Amberlyst 15, Amberlyst 35, Amberlite IR-120, Amberlite 200, Dowex 50, Lewatit SPC 118, Lewatit SPC 108, Bayer K2611, K2621, OC1501, among others), a niobium phosphate mass catalyst (NbPO.sub.3), a Silica-Alumina (SiAl), an acid mixed oxide, titanium and cerium in sulfated zirconia (TiZrSCe), and a prepared zeolite for the production of oligomers, based on H-ZSM5 (Zeolite).

[0107] Tests were performed under various conditions of T, LHSV from 0.5 to 4 h.sup.−1, and typical charge of 50 vol % toluene+50 vol % LPG, with some tests ranging from 20 to 90% aromatic. The base pressure was 60 bar (6 MPa).

[0108] The divinyl-benzene resin (DVB resin) was tested at a temperature of 60 to 140° C. (due to catalyst limitations). The NbPO.sub.3 catalyst was tested from 140 to 250° C. The SiAl catalyst was tested mostly from 200 to 280° C., with some tests up to 380° C. to assess accelerated deactivation. The TiZrSce catalyst has been tested from 140 to 360° C. The zeolitic catalyst was tested from 200 to 320° C.

[0109] We analyzed the conversion of C4= olefins (C4= Olef Conversion, %) versus aromatics yield (Y Arom, %), and results presented in FIG. 2. The objective is the highest possible olefin conversion with the highest yield in aromatics, since the yield in C8= olefins is in the gasoline range, not JET A-1.

[0110] Observing the results, it appears that the divinylbenzene resin (DVB) produces mostly olefins, only increasing in conversions of higher olefins. NbPO.sub.3 had a slightly higher yield in alkyl-aromatics, followed by TiZrSCe. On the other hand, zeolite showed high olefin conversions, but lower yields in alkyl-aromatics. It only showed high yields in alkyl-aromatics with high conversions at higher severities and higher aromatic/olefin ratios than the standard. SiAl silica-alumina combined the desired result of high conversions of olefins with high yields of alkyl-aromatics.

Example 2: Doping with Pt, Pd, Rh and SiAl Catalyst Deactivation Test with and without Metallic Function

[0111] Metals were added to the original SiAl catalyst, by the wet spot impregnation technique, according to WO PCT patent 2001/09628. Drying was carried out in a muffle oven in two steps: 100° C. for 2 h and 140° C. for 2 h (after heating rate of 1° C./min). The calcination was carried out in a muffle furnace in 2 steps, at 300° C. for 2 h and 500° C. for 2 h (at 5° C./min).

[0112] The original catalyst: 0% metals, only silica-alumina, state of the art.

[0113] State of the art catalyst A: 0.2% w/w Pt, 0.6% w/w Pd, prepared with chlorine salts, totaling 0.47% w/w of Cl in the catalyst.

[0114] State of the art catalyst B: 0.2% w/w Pt, 0.6% w/w Pd, prepared with non-chlorine salts in the composition.

[0115] The catalyst of the present invention C: 0.2% w/w Pt, 0.4% w/w Pd, and 0.2% w/w Rh, prepared without chlorine salts.

[0116] Prior to contact with the charge, the catalysts were reduced after loading in the unit at 400° C. for 4 h at 60 bar (6 MPa).

[0117] Tests were carried out to verify the deactivation of the catalysts loaded with LPG+toluene (50/50 vol %), temperature of 230° C., pressure of 60 bar (6 MPa) and LHSV of 2 h.sup.−1.

[0118] FIG. 3 shows the comparison of the first alkylation results (first tests of the catalysts) after the standard reduction procedure, even for the original catalyst, without metals. Catalyst A (PtPd with chlorine) showed greater activity and less deactivation than the original SiAl, but, after oxidative regeneration (Cat A regen Ox), it showed greater deactivation than the original SiAl catalyst. The catalyst B (PtPd without chlorine) showed lower activity than the original SiAl and similar deactivation to the regenerated A. On the other hand, catalyst C showed higher activity than original SiAl, and a lower deactivation tendency than regenerated A, B and original SiAl. The catalyst C of the present invention is more active and with less deactivation than the original support.

[0119] Not only the initial activity and deactivation are important, it is necessary that the hot-stripping procedure with H.sub.2 of the present invention is effective to recover the initial activity of the catalyst and thus allow the catalyst to operate for a long term, avoiding regeneration with oxygen.

[0120] FIG. 4 shows the results for catalyst A, PtPd with chlorine. A first regeneration attempt was carried out, maintaining the flow only of toluene at 300° C. for 8 h (and with enough H.sub.2 excess for 3 times the chemical consumption of hydrogenation). The objective was to verify the impact of the hydrogenation step on the catalyst regeneration. For catalyst A, hot-stripping was performed (passage of H.sub.2 in the reactor, 10 volumes of H.sub.2 per volume of catalyst per minute, for 8 h) at 450° C. after about 350 h, new hot-stripping after 650 h, and a third hot-stripping at 890 h; followed by oxidation regeneration at 980 h (depressurizing, injecting air and N.sub.2 mixture, maintaining 500° C. for 12 h), followed by hot-stripping at 1060 h and 1150 h, as shown in FIG. 4. The results show that the state-of-the-art catalyst with chlorine deactivated more after each hot-stripping step and that the initial activity was not recovered after oxidative regeneration.

[0121] For the state-of-the-art catalyst B, without chlorine, hot striping was performed at 60 h, at 130 h, at 195 h (both at 450° C., the first 2 for 12 h and the third for 24 h). Activity results were similar to the initial test. However, with significantly lower activity than SiAl without metals, and worse than the activity of catalyst A after oxidative regeneration.

[0122] FIG. 5 shows the results of regenerations with H.sub.2 for the catalyst C of the invention, from PtPd+Rh. a first regeneration attempt was performed at 70 h, with toluene and H.sub.2, at 230° C., 60 bar (6 MPa), for 24 h (and with enough excess of H.sub.2 for 3 times the chemical consumption of hydrogenation). Recovery of part of the activity occurred, with an increase over time, which may indicate that the hydrogenation was able to convert part of what deactivated the catalyst, which was removed over time during the alkylation step. This behavior indicates that longer hydrogenation times will likely continue to reactivate the acidic function of the catalyst, even at lower temperatures, compatible with the hydrogenation of aromatics. Test continued until about 145 h, when hot-stripping was performed at 450° C. for 24 h and 60 bar (6 MPa). The same hot stripping was performed at 220 h, 270 h, 330 h. At 400 h, a hot-stripping was performed for 24 h at atmospheric pressure, but with insufficient activity recovery, the hot-stripping was repeated at 425 h, followed by another at 515 h, as shown in FIG. 5.

[0123] Surprisingly, there is no tendency for the deactivation to deteriorate with the continuation of the tests, and the hot-strippings at a temperature higher than the alkylation and pressure equal to the alkylation (450° C. and 60 bar (6 MPa)) were sufficient for the full recovery of the activity of the catalyst, for at least 6 operating cycles. The results indicate that a high number of regenerations with H.sub.2 (hot stripping) can be used before the need for oxidative regeneration.

[0124] The result of the invention of the addition of Rh to the alkylation catalyst containing Pt/Pd, even in a small amount, 0.2% w/w, in the recovery of activity by hot-stripping is surprising and unexpected.

[0125] The alkylation products obtained in the tests were collected for final hydrogenation tests.

Example 3: Use of Catalysts for the Hydrogenation Step

[0126] It was possible to hydrogenate the hydrogenation products obtained in the alkylation step using catalysts B and C.

[0127] The catalysts were compared to a hydrogenation catalyst formulation stream in the diesel range, described in WO PCT 2001/09628, and showed similar results.

Example 4: Formulation with Re

[0128] Preparing catalyst D with Re (0.2% w/w Pt, 0.4% w/w Pd and 0.2% w/w Re) showed similar results to catalyst C in the alkylation, but with lower dealkylation in the hydrogenation step.

[0129] The examples illustrate the claims of the present invention of conversion by alkylation of aromatics and olefinic chains to alkyl-aromatics and alkyl-naphthenics, using a regenerable catalyst, and should not be limiting thereto.