Note: Descriptions are shown in the official language in which they were submitted.
~39~
1 BACKGROU~D OF THE INVENTION AND PRIOR ART
2 Catalytic reforming, or hydroforming, is a well
3 established industrial process employed by the petroleum
4 industry for improving the octane quali~y of naphthas or
straight run gasolines. In reforming, a multi-functional
6 catalyst is employed which conta.ins a metal hydrogenation-
7 dehydrogenation (hydrogen transfer) component, or components,
8 substantiallv atomically dispersed upon the surface of a
9 porous, inorganic oxide support, notably alumina. Noble
metal catalysts, notably of the platinum type, are currently
11 employed, reforming being defined as the total effect of the
12 molecular changes, or hydrocarbon reactions, produced by
13 dehydrogenation of cyclohexanes and dehydrois3merization of
14 alkylcyclopentanes to yield aromatics; dehydrogenation of
paraffins to yield olefins; dehydrocyclization of paraffins
16 and olefins to yield aromatics; isomerization of n-paraffins;
17 isomerization of alkylcyclo-paraffins to yield cyclohexanes;
18 isomerization of substituted aromatics; and hydrocracking of
19 paraffins which produces gas, and inevitably coke, the lat-
ter being deposited on the catalyst.
21 Platinum has been widely commercially used in
22 recent years in the production of reforming catalysts, and
23 platinum-on-alumina catalysts have been commercially em-
24 ployed in refineries for the last few decades. In the last
decade, additional metallic components have been added to
26 platinum as promotors to further improve the activity or
27 selectivity, or both, of the basic platinum catalyst, e.g.,
28 iridium, rhenium, tin, and the like. Some catalysts pos-
29 sess superior activity, or selectivity, or both, as con~
trasted with other catalysts. Platinum-rhenium catalysts by
31 way of example possess admirable selectivity as contrasted
32 with platinum catalysts, selectivity being defined as the
33 ability of the catalyst to produce high yields of C5+ liquid
34 products with concurrent low production of normally gaseous
hydrocarbons, i.e., methane and other gaseous hydrocarbons,
36 and coke.
37 In a conventional process, a series of reactors
38 constitute the heart of the reforming unit. Each reforming
l3
-- 2
1 reactor is generally provided with fixed beds of the cata-
2 lyst which receive upflow or downflow feed, and each is pro-
3 vided with a preheater or interstage heater, because the
4 reactions which take place are endothermic. ~ naphtha feed,
with hydrogen, or recycle gas, is concurrently passed
6 through a preheat furnace and reactor, and then in sequence
7 through subsequent interstage heaters and reactors of the
8 series^ The product from the last reactor is separated into
g a liquid fraction, and a vaporous effluent. The latter is a
gas rich in hydrogen, and usually contains small amounts of
11 normally gaseous hydrocarbons, from which hydrogen is
12 separated for the C5+ liquid product and recycled to the
13 process to minimize coke production.
14 The activity of the catalyst gradually declines
due to the build-up of coke. Coke formation is believed to
16 result from the deposition of coke precursors such as
17 anthracene, coronene, ovalene and other condensed ring
18 aromatic molecules on the catalyst, these polymerizing to
19 form coke. During operation, the temperature of the process
is gradually raised to compensate for the activity loss
21 caused by the coke deposition. Eventually, however, econom-
22 ics dictates the necessity of reactivating the catalyst.
23 Consequently, in all processes of this type the catalyst
24 must necessarily be periodically regenerated by burning off
the coke at controlled conditions.
26 mwo major types of reforming are generally prac-
27 ticed in the multi reactor units, both of which necessitate
28 periodic reactivation of the catalyst, the initial sequence
29 of which requires regeneration, i.e., burning the coke from
the catalyst. Reactivation of the catalyst is then com-
31 pleted in a sequence of steps wherein the agglomerated metal
32 hydrogenation-dehydrogenation components are atomically re-
33 dispersed. In the semi-regènerative process, a process of
34 the first type, the entire unit is operated by gradually and
progressively increasing the temperature to maintain the
36 activity of the catalyst caused by the coke deposition, un-
37 til finally the entire unit is shut down for regeneration,
38 and reactivation, of the catalyst. In the second, or cyclic
~8~3
- 3
1 type of process, the reactors are individually isolated, or
2 in effect swung out of line by various manifolding arrange-
3 ments, motor operated valving and the like. ~he catalyst is
4 regenerated to remove the coke deposits, and then reactivat-
ed while the other reactors of the serles remain on stream.
6 A "swing reactclJ' temporarily replaces a reactor which is
7 removed from the series for regeneration and reactivation
8 of the catalyst, until it is put back in series.
9 Various improvements have been made in these pro-
cesses to improve the per~ormance of reforming catalysts in
11 order to reduce capital investment or improve C5~ liquid
12 yields while improving the octane quality of naphthas and
13 straight run gasolines. Platinum-rhenium catalysts, among
14 the handful of successful con~ercially known catalysts,
mainta n a rank of eminence as regards their se]ectivity;
16 and they have good activity.
17 Variations have been made in the amount, and kind
18 of catalysts charged to the different reforming reactors of
19 a series to modify or change the nature of the product, or
to improve C5~ liquid yield.
21 Different catalysts, ~Tith differing catalytic
22 metal components, have also been used in the different re-
23 actors of the series. U.S. 3~684,692 employs platinum-
24 rhenium catalysts in the various reactors, the lead re-
actors employing a neutral alumina support whereas the tail
26 reactors employ an acidic oxide support for dehydrocycliza-
27 tion. The acidic support contains 90~ alumina and 10~
28 ~-exchanged synthetic faujasite; and, both catalysts are
29 chlorided.
In U.S. 3,660,271 to Keith et al there is des-
31 cribed a process for the catalytic reforming of naphthene
32 and paraffin-containing hydrocarbons for providing a product
33 of improved octane. The first reactor, or reactors, of the
34 series contains a supported platinum-group metal containing
low acidity catalyst which is devoid, or essentially devoid
36 of rhenium, i.e., contains less than about 0.05 wt ~,
37 preferably less than Q.01 wt. % or no detectakle amount of
38 rhenium, which serves to dehydrogenate naphthenes. The
~39~3~3
-- 4
1 tail reactor, or reactors, and preferably the last reactor
2 of the series contains a supported platinum group metal and
3 rhenium containing catalyst of higher acidity which serves
4 to dehydrocyclize paraffins. An e~ample describes a four
reactor system in which a chloxided platinum metal catalyst
6 (0.6~ Pt/0.7~ Cl/A1203) is employed in the first three re-
7 actors of the series, and a chlorided platinum-rhenium
8 catalyst (0.6% Pt/0.6% Re/10.7% Cl/90% A12O3/10~ H-faujasite)
9 is employed in the last reactor of the !series. U.~.
3,705,095 ~o M. H. Dalson et al is quite similar to the
11 Keith et al patent except that the catalysts charged into
12 the several reactors are supported on alumina. In U.S.
13 3,658,691 and U.S. 3,705,094, both to Xeith et al, the
14 rhenium-containing catalyst is employed in the lead reactors
of the series~ In the former there is thus described a pro-
16 cess wherein a platinum-rhenium-chloride/acidic oxide cata-
17 lyst is employed in the initial dehydrogenation reactor, and
18 a platinum-chloride/alumina catalyst is employed in the de-
19 hydrocyclization tail reactor; and in the latter, platinum-
rhenium-chloride/alumina is employed in the first three re-
21 actcrs and a platinum~loride/alumina catalyst is employed
22 in the last reactor of the series. In British 1,470,887 a
23 platinum-rhenium catalyst is also employed in the early
24 stages, and a platinum-iridiur~ catalyst is employed in the
tail reactor. U.S. Patent Nos. ~,167,473; 3,516l924; and
26 4,174,270 are also of interest.
27 l~hereas these variationsl and modifications have
28 generally resulted in improving the process with respect to
29 some selected performance objectivel or another~ it is none-
theless desirable to provide a new and improved process
31 which is capable of achieving highex conversions of the
32 product to C5~ liquid naphthas as contrasted with present
33 reforming operations.
34 This object and others is achieved in accordance
with the present invention, embodying a process for improving
36 the octane quality of a naphtha in a reforming unit comprised
37 of a plurality of serially connected reactors, inclusive of a
38 lead reactor, a tail reactor and optionally one or more reactors
39 intermediate the lead reactor and the tail reactox, each of
38~3
-- 5
1 which contains a platinum-rhenium catalyst, the naphtha flowing in
2 sequence from one reactor of the series to another and contacting
3 the catalyst at reforming conditions in the presence of hydrogen,
4 characterized by maintainin~, in the tail reactor, a catalysk
having a weight ratio of rhenium:platinum of at least about 1.5:1,
6 and in the lead and intermediate reactors a catalyst having a
7 weight ratio of rhenium:platinum of up to abou-t 1:1. The beds
8 o~ catalys~ are contacted with a hydrocarbon or naphtha feed, and
9 hydrogen, at reforming conditions to produce a hydrocarbon, or
naphtha product of improved octane, and the product is withdrawn.
11 In its preferred aspects, the leadin~ reforming
12 zones, or xeactors of the series are provided with platin-
13 um rhenium catalyst wherein the wei~ht ratio of the
14 rhenium:platinum ranges from abou~ 0.1:1 to about 1:1, and
preferably from about 0.3:1 to about 1:1, and the last
16 reforming zone, or reactor of the series is provided with
17 a platinum-rhenium catalyst wherein the weight ratio of
18 therhenium:platinum ranges from about 1.5:1 to about 3:1,
19 and preferably from about 2:1 to about 3:1.
It is known that the amount of coke produced in
21 an operating run increases progressively from a leading
22 reactor to a subsequent, or from the first reactor to the
23 last, or tail reactor of the series as a consequence of
24 the different types of reactions that predominate in the
several different reactors. Thus, in the first reactor
26 of the series the meial site, or hydrogenation-dehydro-
27 genation component of the catalyst, plays a dominant role
28 and the predominant reaction involves the dehydrogenation
29 of naphthenes to aromatics. This reaction proceeds at
relatively low temperature, and the coke formation is
31 relatively low. In the intermediate reactors (usually a
32 second and third reactor), on the other hand, the acid
33 site plays a major role and the isomerization reactions
34 predominate, though additional naphthenes are formed and
these are dehydrogenated to aromatics as in the first
36 reactor. In both of the intermediate reactors the
~L~8~ 3
-- 6 --
1 temperature is maintained higher than in the first reactor,
~ and the ternperature in the third reactor is maintained
3 higher than that of the second reactor of the series.
4 Carbon formation is higher i.n these reactors than in the
first reactor of the series, and coke is higher in the
6 third reactor than in the second reactor of the series.
7 The chief reaction in the last, or tail reactor of the
8 series involves dehydrocycliæation of p~raffins and
9 olefins, and the highest temperature is employed in this
reactor. Coke formation is highest in this reactor, and
11 the reaction is often the most difficult to control. It
12 is also generally known that these increased levels of
13 coke in the sev~ral reactors of the series causes con-
14 siderable deactivation of the catalysts, Whereas the
relationship between coke formation, and rhenium promotion
16 to increase catalyst selectivity is not Xnown with any
17 degree of certainty because of the extreme complexity of
18 these reactions, it is believed that the presence of the
19 rhenium minimizes the adverse consequences of the increased
coke levels, albeit it does not appear to minimize coke
21 formation in any absolute sense. Nonetheless, in accor-
22 dance with this invention, the concentration of the rhenium
23 is increased in those reactors where coke formation is the
24 greatest, but most particularly in the last reactor of the
series. Thus, in one of its forms the catalysts within
26 the series of reactors are progressively staged with
27 respect to the rhenium concentration, the rhenium concen-
28 tration being increased from the first to the last
29 reactor of the series such that the rhenium content of
the platinum-rhenium catalysts is varied significantly to
31 counteract the normal effects of coking,
32 In one of its aspects, optimum utili7ation of
33 rhenium-promoted platinum catalysts is obtained by
34 providing the catalyst of the initial, or first reactor
of the series wi~th rhenium in concentration adequate to
36 provide a weight- ratio of rhenium:platinum ranging from
37 about 0.1:1 to about 0.5:1, preferably from about 0.3:1
38 to about 0.5:1. The catalyst of the intermediate reforming
!
38~3
zones, as represented by the reactors intermediate between the first and last
reactors of the series, are provided whith rhenium in concentration adequate
to provide a wclgllt ratioll of rhenium:platinllm ranging from ~bout O.S:l to
about l:l, preferably above about 0.5:1 to about 0.8:1. The last reactor of
the series is provlded wLth rilellium iu concentration adequate to provide a
weight ratio of rhenlum:platinum from about 1.5:1 to about 3:1, preferably
from about 2:1 to about 3:1. The last reactor of a series, whether the series
contains less than three or more than three reactors, is always provided with
a catalyst which contains a weight ratio of rhenium:platinum of at least 1.5:1
and preferably contalns an atomic ratio of rhenium:platinum ranging from about
2:1 to about 3:1.
The catalyst employed in accordance with this invention ls
necessarily constituted of composite particles which contain, besides a
carrier or support material, a hydrogenation-dehydrogenation component, or
components, a halide component and, preferably, the catalyst is sulfided. The
support material is is constituted of a porous, refractory inorganic oxide,
particularly alumina. The support can contain, e.g., one or more of alumina,
bentonite, clay, diatomaceous earth, zeolite, silica, activated carbon,
magnesia, zirconia, thoria, and the like; though the ~ost preferred support is
alumina to which, if desired, can be added a suitable amount of other
refractory carrier materials such as silica, zirconia, magnesia, titania,
etc., usually in a range of about 1 to 20 percent, based on the welght 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, preferably from about 100 to
about 300 m2/g, a bulk density of about 0.3 to l.O g/ml, preferably about
0.4 to 0.8 g/ml, an average pore volume of about 0.1 to l.l ml/g, preferably
about 0.3 to 0.8 ml/g, and an average pore diameter of about 30 to 300A.
The metal hydrogenation-dehydrogenation component can be composited
with or otherwise intimately associated with the porous inorganic oxide
support or carrier by
6174-1 - 7 ~
8~
1 various techniques known to the art such as ion-exchange,
2 coprecipitation with the alumina in the sol or gel form,
3 and the like. For example, the catalyst composite can be
4 formed by adding together suitable reagents such as a salt
S of platinum and ammonium hydroxide or carbonate, and a salt
6 of aluminum such as aluminum chloride or aluminum sulfate
7 to form aluminum hydroxide. The aluminum hydroxide con-
8 taining the salts of platinum can then be heated, dried,
9 formed into pellets or extruded, and then calcined in
10 nitrogen or other non-agglomerating atmosphere. The metal
11 hydrogenation components can also be added to the catalyst
12 by impregnation, typically via an "incipient wetness"
13 technique which requires a minimum of solution s~ that the
14 total solution is absorbed, initially or after some evapora-
15 tion.
16 It is preferred to de~osit the platinum and
17 rhenium metals, and additional metals used as promoters, if
18 any, on a previously pilled, pelleted, beaded, extruded, or
19 sieved particulate support material by the impregnation
20 method. Pursuant to the impregnation method, porous
21 refractory inorganic oxides in dry or solvated state are
22 contacted, either alone or admixed, or otherwise incorporated
23 with a metal or metals-containing solution,-or solutions,
24 and thereby impregnated by either the "incipient wetness"
25 technique, or a technique embodying absorption from a
26 dilute or concentrated solution, or solutions, with sub-
27 sequent filtration or evaporation to effect total uptake of
28 the metallic components~
29 Platinum in absolute amount, is usually supported
30 on the carrier within the range of from about 0.01 to 3
31 percent, preferably from about 0.05 to 1 percent, based on
32 the weight of the catalyst (dry basis). Rhenium, in
33 absolute amount, is also usùally supported on the carrier
34 in concentration ranging from about 0.1 to about 3 percent,
35 preferably from about 0.5 to about 1 percent, based on the
36 weight of the catalyst (dry basis). The absolute concentra-
37 tion of each, of course, is preselected to provide the
38 desired atomic ratio of rhenium:platinum for a respective
8~8~L~
g
1 reactor of the unit, as heretofore expressed. In the tail
2 reactor, the rhenium is provided in major an~ount relative
3 to the platinum whereas, in contrast, in all other reactors
4 the rhenium is provided in minor amount, or no more than
5 about an equal amount, relative to the platinum, based on
6 the atomic weight of these metals, one with respect to the
7 other. In compositing the metals with the carrier, es-
8 sentially any soluble compound can be used, but a soluble
9 compound which can be easily subjected to thermal decompo-
10 sition and reduction is preferred, for example, inorganic
11 salts s~lch as halide, nitrate, inorganic complex compounds,
12 or organic salts such as the complex salt of acetylacetone,
13 amine salt, and the likeO ~here, e.g., platinum is to
14 be deposited on the carrier, platin~ chloride, platinum
15 nitrate, chloroplatinic acid, ammonium chloroplatinate,
16 potassium chloroplatinate, platinum polyamine, platinum
17 acetylacetonate, and the like, are preferably used. A
18 promoter metal, or metal other than platinum and rhenium,
19 when employed, is added in concentration ranging from
20 about 0.01 to 3 percent, preferably from about 0.05 to
21 about 1 percent, based on the weight of the catalyst.
22 To enhance catalyst performance in reforming
23 operations, it is also required to add a halogen component
24 to the catalysts, fluorine and chlorine being preferred
25 halogen components. The halogen is contained on the
26 catalyst within the range of 0.1 to 3 percent, preferably
27 within the range of about 1 to about 1.5 percent, based on
28 the weight of the catalyst. When using chlorine as a
29 halogen co~ponent, it is added to the catalyst within the
30 range of about 0.2 to 2 percent, preferably within the
31 range of about 1 to 1.5 percent, based on the weight of
32 the catalyst. The introduction o~ halogen into catalyst
33 can be carried out by any mèthod at any time. It can be
34 added to the catalyst during catalyst preparation, for
35 example, prior to, following or simultaneously with the
36 incorporation of the metal hydrogenation-dehydrogenation
37 component, or components. It can also be introduced by
38 contacting a carrier material in a vapor phase ox liquid
-- 10 --
1 phase with a halogen compound such as hydrogen fluoride,
2 hydrogen chloride, ammonium chloride, or the like.
3 The catalyst is dried by heating at a temperature
4 above about 80F, preferably between about 150F and 300F,
in the presence of nitroge~ or oxygen, or both, in an air
6 stream or u~der vacuum. The catalyst is calcined at a
7 temperature between about 500F to 1200F, preferably about
8 500F -to lOOO~F, either in the presence of oxygen in an air
9 stream or in the presence of an inert gas such as nitrogen.
Sulfur is a h.ighly preferred component of the
11 catalysts, the sulfur content of the catalyst generally
12 ranging to about 0.2 percent, prefexabl~ from about O.OS
13 percent to about 0.15 percent, based on the weight of the
14 catalyst (dry basis). The sulfur can be added to the
catalyst by conventional methods, suitabl~ by breakthrough
16 sulfiding of a bed of the catalyst with a sulfur-containing
17 ~aseous stream, e.g., hydrogen sulfide in hydrogen,
18 performed at temperatures ranging from about 350F to about
19 1050F and at pressures ranging from about 1 to about ~0
atmospheres for the time necessary to achieve breakthrough,
21 or the desired sulfur level.
22 The feed or charge stock can be a virgin naphtha,
23 cracked naphtha, a naphtha from a coal liquefaction pro-
24 cess, a Fischer-Tropsch naphtha, or the like. Such feeds
can contain sulfur or nitrogen, or both, at fairly high
26 levels. ~ypical feeds are those hydrocarbons containing
27 from about 5 to 12 carbon atoms, or more preferably from
28 about 6 to about 9 carbon atoms. Naphthas, or petroleum
29 fractions boiling within the range of from about 80~F to
about 450~F, and preferably from about 125F to about
31 375F, contain hydrocarbons of carbon numbers within these
32 ranges. Typical fractions thus usually contain from about
33 15 to about 80 vol.% paraffins, both normal and branched,
34 which fall in the range of about Cs to C12, from about 10
to 80 vol.% of naphthenes falling within the range of from
36 about C6 to C12, and from 5 through 20 vol.%of the
37 desirable aromatics falling within the range of from about
38 C6 to C12-
1 The re~orming runs are initiated by adjusting the
2 hydrogen and feed rates, and the temperature and pressure
3 to operating conditions. The run is continued at optimum
4 reforming conditions by adjustment of the major process
S variables, within the ranges described below:
6 Major Operating Typical Process Preferred Process
7 Varia~les Conditions Conditions
~ . . .
8 Pressure, Psig 50-750 100-300
9 Reactor Temp., F 900-1200 950-1050
10 Recycle Gas Rate,
11 SCF/~ 1000-10,000 1500-3000
12 Feed Rate, W/Hr/W 0.5-10 2.5-5
13 The invention will be more fully understood by
14 reference to the following comparative data illustrating
~5 its more salient features. All parts are given in terms
16 of weight except as otherwise specified.
17 A series of platinum-rhenium catalysts of high
18 rhenium content were prepared for demonstrative purposes
19 from portions of particulate alumina of the type conven-
tionally used in the manufacture of commercial reforming
21 catalysts. These portions of alumina, i.e., 1/16 inch dia-
22 meter extrudates, were calcined for 3 hours at 1000F fol-
23 lowed by equilibration with water vapor for 16 hours.
24 Impregnation of metals upon the supports in each instance
was achieved by adding H2PtC16, HReO4, and HCl in aqueous
26 solution, while carbon dioxide was added as an impregnation
27 aid. After a two hour equilibration, a mixture was
28 filtered, dried, and then placed in a vacuum oven at 150C
29 for a 16 hour period.
Prior to naphtha reforming, the catalyst was
31 heated to 950F in 6% 2 (94% N2), and then soaked in
32 C12/O2 (500 ppm C12, 6% 2~ 5000 ppm H2) for one hour.
33 Following 3 hours in 6% 2 at 950F, the catalyst was
34 cooled to 850F, reduced with 1.5% H2 in N2, and then
presulfided with H2S in this reducing gas to achieve the
36 desired catalyst sulfur level.
37 The platinum-rhenium catalysts employed in other
38 than the tail reactor were conventional, and were obtained
- 12 -
1 already made from a catalyst manufacturer. However, these
2 catalysts are made in a similar manner with the catalysts
3 employed in other than the tail reactor.
4 Inspections on the feeds employed in the tests
are given in Table I.
6 TABLE I
7 Lt. Arabian Persian Gulf
8 Virgin NaphthaPar~fflnic Naphtha
9 API Gravity 59.7 58.9
10 Sulfur, wppm 0.5 0.5
11 Nitrogen, wppm, ~0.1 <0.1
12 Bromine No. cg/g0.12 0.1
13 ASTM Distillation
14 IBPF 180 166
5% 213 203
16 10 219 214
17 20 232 227
18 30 242 239
19 40 255 253
267 269
21 60 278 283
22 70 294 299
23 80 308 315
24 90 324 333
336 346
26 FBP 382 358
27 Demonstration
28 In a first cyclic simulation reforming run (Run
29 1), a high rhenium catalyst containing 0.3% Pt/0u67~
30 Re/1.1% C12/0.15% S for use in the several reactors of a
31 four reactor unit, with all four reactors on-stream was
32 prepared as previous]y described. In a second run (Run
33 2) all of the reactors o~ the series were provided with low
34 rhenium catalysts containing 0.3% Pt/0.3% Re/1.1% C12/0.15%
35 S. The runs were conducted by passing the Light Arabian
36 paraffinic naphtha through the series of reactors at
37 950~F E.I.T., 175 psig, 3000 SCF/B which are the conditions
38 necessary to produce a 102.0 RONC product~ The results
39 given in Table II were obtained, to wit:
- 13 -
1 T~BLE II
2 Catalyst Yield
3 Activity Units C5+ LV%
4 Run 1 (high rhenium) 96.0 69.3
Run 2 (low rhenium) 102.0 72.0
6 These data thus show that the use of the high
7 rhenium catalysts in the several reactors of the series
8 considerably decreased the C5+ liquid yield, and octane
9 number. This is believed due to the "cracking out" of
aromatics precursors in the lead reactors. This conclusion
11 is supported too by the 20 percent increase in light
12 petroleum gases, principally C3 and C4 hydrocarhons, pro
13 duced with the high rhenium catalysts.
14 EXAMPLE 1
1~ A third run (Run 3) was conducted under similax
16 conditions wi~h the same feed except that the three lead
17 reactors were charged with the low rhenium catalysts, and
18 the tail reactor only was charged with the high rhenium
19 catalyst (28 wt.% of the total catalyst charge). The
results, which are compared with the preceding "low
21 rhenium" run, are given in Table III.
22 TABLE_III
23 Catalyst Yield
24 Activity Vnits C5+ LV%
25 Run 2 (low rhenium) 102.0 72.0
26 Run 3 (low rhenium 102.0 72.5
27 lead/high rhenium
28 tail reactor)
29 A C5+ liquid yield credit is thus obtained by
staging the low and high rhenium catalysts as described.
31 The C5~ liquid yield credit i~ further confirmed by the
32 increase in recycle gas hydrogen p~rltv (~Ll%~ for the
33 staged reactor system.
34 EXAMPLE 2
In a fourth run tRun 4), a dry, calcined
36 catalyst containing 0.29% Pt/0.72% Re/1.1% C12/0.14% S
37 was charged to the fourth, or tail reactor of a unit, and
38 the first three reactors were charged with the low rhenium
~ 14 -
1 catalyst. This run was conducted with a more difficult to
2 reform Pe.rsian Gulf Paraffinic naphtha at 950DF F.I.T., 175
3 psig, 3000 SCF/B, at space velocity sufficient to produce a
4 100 RON product. This run in compared to a fifth run (Run
5) conducted at identical conditions with low rhenium
6 catalyst in all four of the reactors, as given in Table IV.
7 TABLE IV
8 Catalyst Yield
9 Activity Vnits C5+ LV%
10 Run 4(low rhenium92.0 75.5
11 lead/high rhenium
12 tail reactor)
13 Run S (low rhenium) 77.0 74.3
14 The improvement in catalyst activity, and yield
is thus manifest. In addition to the improved activity,
16 and high yield advantage utilizing the more difficult
17 feed stock, the run utilizing the high rhenium catalyst in
18 the tail reactor also demonstrates a far greater coke
19 tolerance than the conventional run even at high severity
conditions (972F EIT) in the tail reactor, as shown by
21 the data of Table V.
22 TABLE V
23 H2 Yield Cl-C4
24 Mole % H2 inWt.% on Yield Wt.%
Recycle-Gas Feed_ n Feed
26 Run 4 (low rhenium 77.1 2.31 17.86
27 lead/high rhenium
28 tail reactor)
29 Run 5 (low rhenium) 76.2 2.26 18.82
30 It is apparent that various modifications and
31 changes can be made without departing from the spirit and
32 scope of the present invention, the outstanding feature
33 of which is that the octane quality of various hydrocarbon
34 feedstocks, inclusive particularly of paraffinic feed-
stocks, can be upgraded and improved.
