CA1278544C - Catalytic reforming process - Google Patents

Abstract

ABSTRACT OF THE DISCLOSURE

(In a process wherein, in a series of reforming zones, or reactors, each of which contains a bed, or beds of a sulfur-sensitive polymetallic platinum-containing catalyst, the beds of catalyst are contacted with a hydrocarbon or naphtha feed, and hydrogen, at reforming conditions to produce a hydrocarbon, or naphtha product of improved octane, the improvement wherein, at start-up, sulfur is added to the tail reactor of the series, and excluded from the lead reactor. Increased hydrogen purity, aromatics, and C5+ liquid yields are obtained, and there is less gas make.

Description

127854~

~ACKGROUND OF THE INVENTION AND PRIOR ART
I. Field of the Invention Thi3 invention rela~e3 to a process for re~orming with sulfur-sensitive, polymetallic platinum-containing catalysts wherein sulfur is added to an intermediate or final reactor, or reaction zone, or zones, and excluded from the initial reactor, or reaction zone, of the series.
II. Background _nd_Problems Catalytic reforming, or hydroforming. is a well establi3hed industrial process employed by the petroleum industry for improving the octane quality of naphthas or straight run gasolines. In reforming, a multi-functional catalyst is employed which contains an acid component and a metal hydrogenation-dehydrogenation (hydrogen transfer) com-ponent, or components, substantially atomically dispersed upon the 3ur~ace of a porous-, inorganic oxide support, notably alumina. Nobie metal catalysts, notably of the platinum type, are currently employed as metal hydrogena-tion-dehydrogenation components, reforming being defined as the total effect of the molecular changes, or hydrocarbon reactions, produced by dehydrogenation of cyclohexanes and dehydroisomerization of alkylcyclopentanes to yield aromatics; dehydrogenation of paraffins to yield olefins;
dehydrocyclization of paraffins and olefins to yield aromatics; isomerization of n-paraffins; isomerization of alkylcycloparaffins to yield cyclohexanes; isomerization of s-ubstituted aromatics; and hydrocracking of paraffins which produces gas, and inevitably coke, the latter being progres-sively deposited on the catalyst as reforming is continued.
In a reforming operation, one or a series of reac-tors, or a series of reaction zones, are employed. Typi-cally, a series of reactors are employed, e.g., three or four reactors, these constituting the heart of the reformin3 unit. Each reforming reactor is generally provided with a fixed bed, or beds, of the catalyst which receive downflow ~;~'7~5~

feed, and each i~ provided with a preheater or interstags heater. A naphtha feed, with hydrogen, or recycle hydrogen gas, is concurrently pa~sed through a preheat furnace and reactor, and then in sequence through sub~equent interstage heaters and reactors of the series. The product from the last reactor is separated into a liquid fraction, and a vaporous effluent. The former is recovered as a C5+ liquid product. The latter is a gas rich in hydrogen, and u~ually contains ~mall amounts of normally gaseous hydrocarbons, from which hydrogen is separated and recycled as "recycle sas" to the proces~ to minimize coke production. In conven-~ onal operations, the recycle gas, which generally contains moisture and hydrogen sulfide impurities, i3 passed through a recycle gas drier which removes much of the moisture and hydrogen sulfide prior to the introduction of the recycle gas into the first reactor of the series.
The ~um-total of the reforming reactions, supra, occurs as a continuum between the fir~t and last reactor o~
the series, i.e., as the feed enter~ and passes over the ~irst fixed catalyst bed of the first reactor and exits from the last fixed catalyst bed of the last reactor of the series. The reactions which predominate between the several reactors differ dependent principally upon the nature of the feed, and the temperature employed within the individual reactors. In the initial reaction zone, or first reactor, which is maintained at a relatively low temperature, the primary reaction involves the dehydrogenation of naphthenes to produce aromatic~. The isomerization of naphthenes, notably C5 and C6 naphthene , also occurs to a considerable extent. Most of the other reforming reactions also occur, but only to a lesser, or smaller extent. There is relative-ly little hydrocracking, and very little olefin or paraf.in dehydrocyclization occurs in the first reactor. Naphthene dehydrogenation is an endothermic reaction, and consequently the reactions in the first reactor are extremely endo-thermic, generally accounting for as much as 2/5 to 3/5 o~
the total observed temperature difference (~T) across the 12'7~3~;44 several catalyst beds contaLned in the several reactors of the series. Within the intermediate reactor(s), or reaction zone(s), the temperature 19 maintained somewhat higher than in the first, or lead reactor of the series, and it is believed that the primary reactions in the intermediate reactor, or reactors, involve the i30merization of naph-thenes and paraffin~. Where, e.g., there are two reactors disposed between the first and last reactor of the series, it is believed that the principal reaction involves the isomerization of naphthenes, normal paraf~ins and isopara~-fins. Some dehydrogenation of naphthenes may, and usually does occur, at least within the first of the intermediate reactors. There is usually some hydrocracking, at ieast more than in the lead reactor of the series, and there is more olefin and paraffin dehydrocyclization. The net effect of the reaction~ which occur in this reactor are endother-mic, and though the temperature drop between the feed inlet and feed outlet i9 not a3 large as in that of the initial reactor (even though the second reactor generally contains a larger catalyst charge than the initial reactor), it is nonetheless considerable. The third reactor of the series, or second intermediate reactor, is generally operated at a somewhat higher temperature than the second reactor of the series. It i3 believed that the naphthene and paraffin isomerization reactions continue as the primary reaction in this reactor, but there is very little naphthene dehydro-genation. There is a further increase in paraffin dehydro-cyclization, and more hydrocracking. The net effect of the reactions which occur in this reactor is also endothermic, though the temperature drop between the feed inlet and feed outlet is smaller than in the first two reactors. In the final reaction zone, or final reactor, which is typically operated at the highest temperature of the series, it is believed that paraffin dehydrocyclization, particularly the dehydrocyclization of the short chain, notably C6 and C7 par~ffins, is the primary reaction. The isomerization reac-tions continue, and there is more hydrocracking and coke formation in this reactor than in any o~ the other reactors of the series. The net e~fect of the reactions which occur in this reactor is al90 generally endothermic.
Platinum is widely commercially used in the pro-duction of reforming catalysts, and platinum-on-alumina catalysts have been commercially employed in refineries ~or the last ~ew decades. In more recent years polymetallic cata~yats have been used. These are catalySts wherein addi-tional metallic components have been added to platLnum as promoters to further improve the activity or selectivity, or both, of the basic platinum catalyst, e.g., iridium, rhenium, tln, and the like. Such catalysts po~sess superior activity, or selectivity, ar both, as contrasted with the basic platinum catalyst. Platinum-rhenium catalyst3 by way of example possess admirable selectivity as contrasted with platinum catalyst~, selectivity being defined as the ability of the catalyst to produce high yields of C5 liquid pro-ducts with concurrant low production of normally gaseous hydrocarbons, i.e., methane and other gaseous hydrocarbons, and coke.
The activity of the catalyst gradually declines due to the build-up of coke. Co~e formation is believed to result from the deposition of coke precursors such as anthracene, coronene, ovalene, and other condensed ring aromatic molecules on the catalyst, these polymerizing to form coke. During operation, the temperature o~ the process is gradually raised to compensate ~or the activity 1099 caused by the coke deposition. Eventually, however, temper-ature increages cannot compensate for tne loss in catalytic activity and hence it becomes necessary to reactivate the catalyst. Consequently, in all processes of thi~ type the catalyst must necessarily be periodically regenerated by burning off the coke at controlled conditions.
Two major types of reforming are generally prac-ticed in the multi reactor units, both o~ which necessitate periodic reactivation of the catalyst, the initial sequence of which requires regeneration, i.e., burning the coke from 1278S~4 the catalyst. Reactivation of the catalyst Ls then com-pleted in a sequence of steps wherein the agglomerated metal hydrogenation-dehydrogenation component~ are atomLcally redispersed. In the semi-regenerative proce3s, a proceas of the rirst type, the entire unit is operated by gradually and progressively Lncreasing the temperature to maintain the actlvity of the catalyst caused by the coke deposition, until finally the entire unit is shut down for regeneration, and reactivation, of the cataly3t. In the second, or cyclic type of proce~s, the reactors are individually isolated, or in effect ~wung out of line by various manifolding arrange-ments, motor operated valving and the like. The catalyst is regenerated to remove the coke depo~its, and then reacti-vated while the other reactors of the series remain on stream. A "3wing reactor" temporarily replaces a reactor which is removed from the series for regeneration and reactivation of the catalyst, until it is put back in series.
Change in the total, or overall, aT is a good indication of changing performance in the reactors during an operating run, and correlates well with the ability of the reaction system to produce reformate octane value; which normally decrease~ throughout the run. In cyclic opera-tions, in particular, the decline in the temperature drop across a catalyst bed is sometimes used as a criterion for selecting the next reactor candidate for regeneration of its catalyst charge.
Essentially all petroleum naphtha or synthetically derived naphtha feeds contain sulfur, a well known catalyst poison which can gradually accumulate upon and poison reforming catalysts. Most of the sulfur, because of this adverse effect, is removed from feed naphthas, e.g., by hydrofining and subsequent contact with guard beds packed or filled with sulfur ad30rbents. The polymetallic refor~ing catalysts are particularly sulfur-sensitive, and particular-ly susceptible to sulfur poisoning. The presence of even small and virtually infinitesimal amount3 of sulfur in ~2t7~S4~

either the naphtha feed or hydrogen recycle gas, or both, lt has been observed, adversely affects 5he process, and per-formance of the catalysts. Wherea3 various improvements have been made in adsorbents, and in the operation of guar~
beds to eliminate sulfur ~rom the naphtha feed and hydrogen recycle gas, the complete elimination of sulfur from the naphtha does not appear practical, if indeed poqsible, and sulfur inevita~ly appear3 in the process. The effect of sulfur in the naphtha, even in concentration ranging only a few parts per million is that, in the overall reforming operation, the yield of hydrogen, aromatics, and C5l liquid yield decreases as sulfur builds up and increases in the system, there is an increase in the rate of cataly3t deacti-vation, and in the total production of C1-C4 light gases.
III. Objects It iq, accordingly, the primary objective of this invention to provide a new and improved proces_ useful in the operation of reforming unit3 which employ highly active sulfur-sensitive polymetallic platinum-containin~ catalyst to produce high octane gasolines.
A specific object is to provide a novel process as characterized, but particularly a star~-up procedure which will provide improved hydrogen purity and aromatics produc-tion, increased C5~ liquid yield, and decreased C4- light gas production in the operation of a reforming unit.
I~. The Invention Theqe objects and others are achieved in accord-ance with the present invention, embodying a process whereln, in a series of reforming zones, or reactors, each of which contains a bed, or beds of catalyst, the catalyst in each of which i9 constituted of a sulfur-3ensitive, polymetallic platinum-containing catalyst which contains little or no coke, and naphtha, and hydrogen, are introduced into the lead reactor, pa3sed in series from one reactor to another, and reacted at reforming conditions, sulfur is introduced into the final or tail reactor of the series to provide and maintain sulfur in concentration within the 8~4~

naphtha to said final reactor ranging from about 0.5 part3 per milllon, based on the weight of the naphtha ~eed (wppm), to about 20 wppm, preferably from about 0.5 wppm to about ô
wppm, while excluding sul~ur from the laad reactor of the series. The sulfur can be added to an intermediate reactor Or the series, or added directly to the ~inal reactor, pref-erably the latter. It ha~ been ~ound that the introduction, or addition of qulfur to the final reactor, or reaction zone of the series, during ~tart-up, or that early portion of the operating cycle when the cataly~t contains lLttle or no coke, relative to the total operating cycle throughout which coke gradually, and progressively builds up and accumulates on the catalyst, provide~ improved hydrogen, aromatics, and C5~ liquid yields, and reduced C1-C4 light gas make.
In accordance with thi~ invention, sulfur is excluded, or its presence minimized, during start-up with a fresh or regenerated catalyst from the lead reactor of the series wherein naphthene dehydrogenation is the predominant reaction. A higher level of sul~ur is maintained during this period in the final reactor of the series wherein paraffin and olefin dehydrocyclization are the predominant reactions. In carrying out such operation, sulfur is injected during this period into an intermediate or the tail reactor, preferably the latter, as hydrogen sulfide, or compound decomposable in 3itu to form hydrogen sulfide, ir.
amount suf~icient to provide sulfur in concentration ranging from about 0.5 wppm to about 20 wppm, prererably 0.5 wppm to about 8 wppm, based on naphtha feed. Sulfur is then removed ~rom the product hydrogen gas from the last reac~or of the ~erLes, suitably by passage o~ the gas through a sulfur adsorbent, and the gas is recycled. The gas i~ recycled to the lead reactor where it is added with fre~h essentially sulfur-free naphtha. The essentially sulfur-free naphtha feed, and hydrogen, entering the lead reactor will provide a concentration of less than 0.5 wppm within the naphtha feed to the first reactor, and preferably no more than about 0.1 wppm sul~ur. It is found that hydrogen, aromatics and C5~

~2~85~

liquid yields are improved, and reduced Cl-C4 light sas ma~e up to such point in time that the catalyst in the final reactor, or reaction zone, contains no more than about 10 percent coke, based on the total weight of the catalyst, deposited thereon; and preferably no more than about 5 per-cent coke deposited upon the catalyst. This period corre-sponds generally from about 5 to about 60 percent of the operating cycle, and preferably from about 5 percent to about 30 percent o~ the total operating cycle which begins at start-up, or when the catalyst i3 first placed on-oil.
The catalyst employed in accordance with this invention is necessarily constituted of composite particles which contain, besideY a carrier or support material, a hydrogenation-dehydrogenation component, or components, a halide component and, pre~erably, the catalyst is sulfid-ed. The support material i~ constituted of a porous, refractory inorganic oxide, particularly alumina. The sup-port can contain, e.g., one or more of alumina, bentonLte, clay, diatomaceous earth, zeolite, silica, activated carbon, magnesia, zirconia, thoria, and the like; though the most preferred support is alumina to which, if desired, can be added a suitable amount of other refractory carrier mate-rials such as silica, zirconia, magnesia, titania, etc., usually in a range of about 1 to 20 percent, based on the weight of the support. A preferred support for the practice of the present invention is one having a surface area of more than 50 m2/g, pre~erably from about 100 to about 300 m2~g, a bulk density of about 0.3 to 1.0 g/ml, preferably about 0.4 to 0.8 g/ml, an average pore volume of about 0.2 to 1.1 ml/g, preferably about 0.3 to 0.8 ml/g, and an average pore diameter of about 30 to 300~.
The metal hydrogenation-dehydrogenation component, or components, includes platinum, and one or more of iridium, rhenium, palladium, rhodium, tin, and tungsten.
Preferably, the hydrogenation-dehydrogenation component, or components, are platinum and iridium or rhenium, or platinum and both iridium and rhenium. The hydrogenation-dehydroge-nation component, or components, can be composited with orotherwise intimately associated with the porous inorganic oxide support or carrier by variou~ techniqueq known to the art such as ion-exchange, coprecipitation with the alumina in the 901 or gel ~orm, and the like. For example, the catalyst composite can be formed by adding together suitable reagents such as a salt o~ platinum and a salt of rhenium and ammonium hydroxide or carbonate, and a 3alt of aluminum such as aluminum chloride or aluminum sulfate to form aluminum hydroxide. The aluminum hydroxide containing the salts of platinum and rhenium can then be heated, dried, formed into pellets or extruded, and then calcined in nitrogen or other nonagglomerating atmosphere. The metal hydrogenation components can also be added to the catalyst by impregnation, typically via an "incipient wetness" tech-nique which requires a minimum of solution 90 that the total solution is absorbed, initially or after some evaporation.
It is preferred to deposit the platinum and iridium or rhenium metals, or both, and additional metals used as promoters, if any, on a previously pilled, pelleted, beaded, extruded, or sieved particulate support material by the impregnation method. Pursuant to the impregnation method, porous refractory inorganic oxides in dry or solvated state are contacted, either alone or admixed, or otherwise incorporated with a metal or metals-containing solution, or solution~, and thereby impregnated by either the "incipient wetness" technique, or a technique embodying absorption ~rom a dilute or concentrated solution, or 301u-tions, with subsequent filtration or evaporation to effect total uptake of the metallic components.
Platinum in absolute amount is usually supported on the carrier within the range of from about 0.05 to 3 percent, preferably from about 0.2 to 1 percent, based on the weight of the catalyst (dry basis). Rhenium, in abso-lute amount, is also usually supported on the carrier in concentration ranging from about 0.05 to about 3 percent, preferably from about 0.3 to about 1 percent, based on the 1278~

-- 1 o -weight of the catalyst (dry basi3). Iridium, or metal other than platinum and rhenium, when employed, is alao added in concentration ranglng from about 0.05 to about 3 percent, preferably ~rom about 0.2 to about 1 percent, based on the weight of the catalyst (dry basis). The absolute concentra-tion of each, of course, is preselected to provide the desired ratio of rhenium:platinum for a respective reactor of the unit, a~ heretofore expre~sed.
In compositing the metal~ with ehe carrier, e3sen-tially any soluble compound can be used, but a soluble com-pound which can be easily ~ubjected to thermal decomposition and reduction is preferred, for example, inorganic salts such as halide, nitrate, inorganic complex compounds, or organic salts such as the complex salt of acetylacetone, amine salt, and the like. When, e.g., platinum i9 deposited on the carrier, platinum chloride, platinum nitrate, chloro-platinic acid, ammonium chloroplatinate, potassium chloroplatinate, platinum polyamine, platinum acetylacetonate, and the like, are preferably used.
To enhance catalyst performance in reforming operations, it is also required to add a halogen component to the catalysts, fluorine and chlorine being preferred halogen components. The halogen is contained on the cata-lyst within the range of 0.1 to 3 percent, preferably within the range of about 0.3 to about 1.5 percent, based on the weight of the catalyst. When using chlorine as halogen com-ponent, it is added to the catalyst within the range of about 0.2 to 2 percent, preferably within the range of about 0.5 to 1.5 percent, based on the weight of the catalyst.
The introduction of halogen into catalyst can be carried out by any method at any time. It can be added to the catalyst during catalyst preparation, ~or example, prior to, follow-ing or simultaneously with the incorporation of the metal hydrogenation-dehydrogenation component, or components. It can also be introduced by contacting a carrier material in a vapor phase or liquid phase with a halogen compound such as hydrogen fluoride, hydrogen chloride, carbon tstrachloride, or the like.

12t78~4 The catalyst i9 dried by heating at a temperature above about 80F, preferably between about 150~F and 300F, in the presence of nitrogen or o~cygen, or both, in an air stream or under vacuum. The catalyst is calcined at a tem-perature between about 500F to 1200F, preIerably about 500F to 1000F, either in the presence of' oxygen in an air stream or in the presence of an inert gas ~uch a3 nitro&en.
Sulfur i~ a highly preferred component of the catalysts, the sulfur content of the catalyst senerally ranging to abou'c 0.2 percent, preferably from about 0.03 percent to about 0.15 percent, based on the weight of the catalyst (dry basis). A fresh reforming cataly~t is generally ~ulfided prior to its being placed on-oil, and 3ince sul~ur i~ lost during reforming, the catalyst i~ again sulf-ided during catalyst regeneration. The sulf'ur can be added to the catalyst by conventional methods, 3uitably by breakthrough sulf'iding of a bed of the cataly3t with a sulrur-containing gaseous stream, e.g., hydrogen sulfide in hydrogen, performed at temperatures ranging from about 350F
to about 1050F and at pre~sures ranging from about 1 to about 40 atmo3pheres for the time necessary to achieve breakthrough, or the desired sulf ur level.
The feed or charge stock can be a virgin naphtha cracked naphtha, a naphtha f`rom a coal liquefaction proces~, a Fischer-Tropsch naphtha, or the like. Typical feeds are those hydrocarbons containing from about 5 to 12 carbon atom3, or more preferably from about 6 to about 9 carbon atoms. Naphthas, or petroleum fractions boiling within the r~nge of from about 80F to about 450F, and preferably from about 125F to about 375F, contain hydrocarbons of carbon numbers within these ranges. Typical fractions thus usually contain from about 15 to about 80 vol. ,~, paraffin3, both normal and branched, which fall in the range of about C5 to C~2, from about 10 to 80 vol. % of naphthenes falling within the range of from about C6 to C12, and from 5 through 20 vol. ~ of the desirable aromatics falling within the range of from about C6 to C12.

~2'7~S~

The rerorm$ng run~ are inltlated by adJustlng the hydrogen and feed rate3, and the temperature and pressure to operatlng condltion~. The run $s contlnued at optimum re~ormlng condltlons by ad~u-~tment o~ the ma~or process variables, with$n the ranges de~crlbed below:
MaJor Operating Typical Process Preferred Process Variables Conditions Conditions Pressure, psig 50-750 100-500 Reactor Temp., F 800-1200 850-1000 Recycle Gas Rate, SCF/B 1000-10,000 1500-5000 Feed Rate, W/Hr/W 0.5-10 1-5 The invention will be ~ore rully understood by re~erence to the following comparative data illustratin~ its more salient reatures. All parts are given in terms of weight except as otherwi~e specified.
A rhenium promoted platLnum catalyst (0.3 wt. ~
Pt/0.3 wt. % Re) obtained from a catalyst manufacturer was employed in the several reactors of the reforming units used in conducting the following runs. The catalyst was reduced and pre ulrided by contact with an admixture of 300-500 vppm (ppm by volume) H2S in hydrogen at 700-8500F until H2S
breakthrough at the reactor outlet, a treatment suf~icient to deposit 0.08-0.10 wt. S sul~ur on the cataly~t initially charged to the unit.
In~pections on the petroleum virgin naphtha ~eedstock used in the te~t for conducting the run described under Example 1, and the naphtha feedstock employed for conductlng the runs described under Examples 2, 3 and 4 are g$ven in the table i~mediately below.
Figures 1 and 2 of the attached drawings respectively illustrate the results of the pilot plant run in Example 2 and the maintenance of the hydrogen yield after increase in the sulfur level during such run.

~78544 Table .

Example 1 Examples 2,3 & 4 Feedstock Feedstock API Gravity 56.2 59.7 Average Mol. Wt. 108 108 Nitrogen, wppm <1 <1 ASTM DistLllation IBPF 180+2 181 +2 309+5 328 F9P 323+10 340+10 .

The following demonstration run, Example 1, illu-3trates the detrimental effects of feed-borne ~ulfur on catalyst in the lead reactor position and net overall re-forming unit performance.

A commercial ref'orming unit with three onstream reactors in serie~, each containing a platinum-rhenium catalyst (0.3 wt. ~ Pti3.3 wt. % Re) in which less than eleven percent of the total onstream catalyst charge was contained in the first of the three reactors, was employed to process eighteen thousand barrels per day of naphtha of average molecular weight ~ 108 and specific gravity (60/600F) ~ 0.754. During the run, after continuous opera-tion to a point at which the catalyst in the tail reactor was estimated to have accumulated from 7 to 8 weight percent coke, the feed sulfur level was increased from an initial level of less than 0.3 wppm sulfur in feed to 1.0 wppm sulfur in feed. Although operating at the lowest average temperature of the three reactors, and containing the smallest amount of catalyst, performance of the initial reactor, or reactor in which the incoming hydrocarbon feed-stock is first contacted with catalyst, was observed to undergo a decrease in the temperature drop between the 1'~78544 reactor Lnlet and outlet, indicating a change in the nature or extent Or reactions taking place wlthin the catalyst bed. Within three days Or the increase in sulfur level, temperature drop of the lead reactor was observed to decrease from 97F to 77~F while average temperature in the catalyst bed increased 10F. Thi~, as would be expected, resulted in an increased rate of catalyst deactivation, which i3 highly temperature dependent. Temperature differ-ences between the inlet and outlet Or the second and third reactors during this time remained essentially unchanged at 84 + 1F and 16 ~ 1F, respectively. Overall unit perfor-mance, as indicated by the total temperature drop acro3s all reactors, deteriorated from 197F temperature drop to 177F
temperature drop. This apparent decrease in unit perfor-mance was corroborated by an observed drop in reforming unit hydrogen yield from 1.70 weight percent Or feed before the sulrur increase to 1.60 weight percent hydrogen yield a~ter-ward, and an approximate 10 percent decrease in unit activ-ity, or a drop in reformate octane of 0.8 RONC at the con-ditions at which the unit was operating.
The following Example 2 substantiates the large sensitivity of overall unit performance to deteriorations in the lead reactor catalyst, showing that the harmful effects of feed-borne sulfur can be prevented if damage to the cata-lyst in the lead reactor position is avoided.

A first pilot plant run was conducted in which a coke-free catalyst of the same composition employed in the p-receding demonstration was contained in a re~orming unit employing three onstream reactors in series in the same proportions as described in the preceding demonstration.
Process conditions of the pilot plant operation were the same as those of the preceding run, 360 + lO psig operating pressure and 3200 + 200 SCF/B recycle gas ratq. The reforming unit in which this evaluation was made permitted adiabatic operation in each of the individual reactors, or the use of an adjustable external heat control means such 1;~7~3544 that a temperature profile could be lmposed. After startup the adiabatic temperature profiles were permitted to develop in all reactors whlle the sulrur level Or the incoming naphtha reedstock was maintained at a constant level of 0.3 wppm. Arter eignt days, at which time the catalyst in all reactors had accumulated less than 4 weight percent coke, as determined from previous tests conducted at the same condi-tions, reedstock sulrur level was increased to 1.5 wppm. To examine the er~ects o~ thi~ change without the perturba-tional ef~ects Or increased averaga temperature in the rirst reactor catalyst bed, the adiabatic temperature drop which exlsted in the rirst reactor cataly3t bed before the ~ulrur level increase was maintained after the increase by means cr an ad~ustable external heat control. This prevented an increased average bed temperature and accelerated rate Or deactivation in the lead reactor, while permitting the changed sulrur level to manirest its er~ects on the catalyst charge contained in the second and third onstream reactors.
- The results of the run are graphically depicted as circular points by rererence to Figure 1. As in the preced-lng demonstration, there was no observed change in the temperature drop between the inlet and outlet Or the second and third reactors. Unlike the preceding demonstration, however, there was a substantial improvement in the purity Or the hydrogen-containing recycle gas. In less than twenty-four hours after the sulfur level increase, the hydrogen content of gas rrom the reactor erfluent product separator vessel, maintained at constant temperature and pressure, was observed to increase by 2 volume percent. The yield of methane, the principal impurity which remains in the excess separator drum gas after downstream puri~ication ror use in other hydrocarbon processes in an integrated refinery, decreased by roughly one-sixth, as also shown in Figure 1. There was no loss in hydrogen yield arter the sulfur level increase. This is illustrated by the data shown in Figure 2. That the additional sulrur added to the system was present throughout both the second and third - l6 -reactors, and not, in view o~ the small total mass involved, merely adsorbed on a portion of the catalyst bed, was con-~irmed by careful measurements of the hydrogen sul~ide con-tent of the separator drum overhead gas.
The following Example 3, when compared with the preceding examples, 3how3 that the presence of feed-borne sulfur early in the unit operatlng cycle, when the catalyst contains no more than about lO weight percent carbon, can in ~act be advantageous when damage to the lead position cata-lyst is avoided.

A 3econd pilot plant run identical to that de-scribed in Example 2 was carried out as in the first pilot plant run, the only difference between the two runs being that the level of sulfur in this second, 3imultaneous control test was maintained at 0.3 wppm throughout. The results obtained Ln thi~ second pilot plant run are ~uperLmposed on the graphs depicted in Figures 1 and 2 as dashed lines, and reference is made to these figures. A
comparison of results from these two evaluations reveals that the incremental sulfur added to the second and third reactors of Example 2 while preventing damage to catalyst in the lead reactor is, in fact, advantageous. Figure l illustrates the higher hydrogen purity and lower methane yield of the performance obtained in the first pilot plant run. These benefltq, as shown by reference to Figure 2, have been realized without any disadvantageous 1099 of hydrogen yield. The overall unit actirities, or ability to produce high-octane reformate at a given set of conditions, of Examples 2 and 3 were equal both at the time o~
incremental sulfur addition in Example 2 and for at least ten days afterward.
The following Example 4 confirms that these benefits are enjoyed during the early portions of the operating cycle between regenerations, up to the time that the catalyst contains about lO weight percent carbon.

~'~78~i44 The first and second pilot plant runs (Example3 2 and 3) were continued past the times graphically depicted .n the figures. The inlet temperatures Or the reactors in each evaluation were increased, as necessary, to compensate for catalyst deactivation and maintain constant reformate pro-duct octane value. The temperature drop across the first reactor cataly3t beds was kept equal to ensure that any observed differences between the two run~ would be manifes-tations of the effects of sulfur addition on the second and third onstream reactors. It was found that the performance advantages obtained in the first p$10t plant run did not extend much beyond the times shown in the illustratLng figures. After about twenty days on oil, a slight activity difference between the two pilot plant runs began to be observed. This difference grew, of course, as temperatures were increased and the rates of deactivation accelerated.
After about forty days on oil, the hydrogen purities of separator drum overhead gas were the 3ame. After about fifty-~ive days on oil, a difference in the weight percent hydrogen yields of the two runs began to develop. To provide assurances of data integrity and userul extended performance results, the runs were continued until catalyst activity had in each case declined to one-fourth of it3 initial, start-of-run value. The runs begun as described in Examples 2 and 3 were terminated after about three and slightly greater than five months, respectively. From the multitude of measurements made over the course of these evaluations, it was found that the benefits obtained in the first pilot plant run are enjoyed during the early portions of the operating cycle, when the reforming catalyst contains up to about tO weight percent carbon, or coke.

Claims (12)

1. In a process for improving the octane quality of a naphtha in a reforming unit comprised of a plurality of serially connected reactors, inclusive of a lead reactor and one or more subsequent reactors, each of which contains a sulfur-sensitive polymetallic platinum-containing catalyst which contains sulfur in concentration ranging from 0.03 to about 0.2 percent, based on the weight of the catalyst, and coke in concentration ranging up to about 10 percent, based on the weight of the catalyst, naphtha and hydrogen being introduced into the lead reactor and flowing in sequence from one reactor to the next subsequent reactor of the series contacting the catalyst at reforming conditions, the improvement comprising, introducing sulfur to the tail reactor of the series and maintaining therein, as a component of the naphtha, sulfur in concentration ranging from about 0.5 wppm to 20 wppm, based on the weight of the naphtha feed, while maintaining the concentration of sulfur in the feed naphtha entering the lead reactor below 0.5 wppm.

2. The process of claim 1 wherein the concentration of sulfur maintained in the naphtha feed to the tail reactor ranges from about 0.5 wppm to about 8 wppm.

3. The process of claim 1 wherein the catalyst of the tail reactor contains from about 0.05 to about 3 weight percent platinum and about 0.05 to about 3 weight percent rhenium.

4. The process of claim 3 wherein the catalyst of the tail reactor contains from about 0.2 to about 1 weight percent platinum and from about 0.3 to about 1 weight percent rhenium.

5. The process of claim 1 wherein the catalyst of the tail reactor contains from about 0.05 to about 3 weight percent platinum and from about 0.05 to about 3 weight percent iridium.

6. The process of claim 5 wherein the catalyst of the tail reactor contains from about 0.2 to about 1 weight percent platinum and from about 0.2 to about 1 weigh. percent iridium.

7. The process of claim 1 wherein the catalyst of the tail reactor contains platinum, rhenium, and iridium.

8. The process of claim 3 wherein the catalyst of the tail reactor contains from about 0.1 to about 3 weight percent halogen.

9. The process of claim 3 wherein the catalyst of the tail reactor contains from about 0.5 to about 1.5 weight percent halogen.

10. The process of claim 3 wherein the catalyst of the tail reactor is sulfided, and contains from about 0.03 to about 0.15 weight percent sulfur.

11. The process of claim 1 wherein the naphtha feed to the tail reactor contains from about 0.5 wppm to about 8 wppm sulfur, and the naphtha feed to the lead reactor contains no more than about 0.1 wppm sulfur.

12. The process of claim 1 wherein the sulfur component on the catalyst within the several reactors of the series ranges from about 0.03 percent to about 0.15 percent.