Moving bed reforming process without heating between the combined feed exchanger and the lead reactor

This invention is a reforming process that employs at least two moving bed reaction zones. One of the reaction zones, called the lead reaction zone, passes catalyst particles and a hydrocarbon-containing effluent to the other reaction zone, which operates at an inlet temperature that is at least 60.degree. F. (33.degree. C.) hotter than the lead reaction zone. In a preferred embodiment, this invention employs no heating between the combined feed exchanger and the lead reaction zone. This invention is particularly applicable to reforming processes that employ continuous regeneration sections.

FIELD OF THE INVENTION 
The field of this invention is the reforming of hydrocarbons. 
BACKGROUND OF THE INVENTION 
Catalytic reforming is a well-established hydrocarbon conversion process 
employed in the petroleum refining industry for improving the octane 
quality of hydrocarbon feedstocks, the primary product of reforming being 
motor gasoline. The art of catalytic reforming is well known and does not 
require detailed description herein. 
Briefly, in catalytic reforming, a feedstock is admixed with a recycle 
stream comprising hydrogen to form what is commonly referred to as a 
combined feed stream, and the combined feed stream is contacted with 
catalyst in a reaction zone. The usual feedstock for catalytic reforming 
is a petroleum fraction known as naphtha and having an initial boiling 
point of about 180.degree. F. (82.degree. C.) and an end boiling point of 
about 400.degree. F. (203.degree. C.). The catalytic reforming process is 
particularly applicable to the treatment of straight run naphthas 
comprised of relatively large concentrations of naphthenic and 
substantially straight chain paraffinic hydrocarbons, which are subject to 
aromatization through dehydrogenation and/or cyclization reactions. 
Reforming may be defined as the total effect 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 substituted aromatics, and 
hydrocracking of paraffins. Further information on reforming processes may 
be found in, for example, U.S. Pat. No. 4,119,526 (Peters et al.); 
4,409,095 (Peters); and 4,440,626 (Winter et al.). 
A catalytic reforming reaction is normally effected in the presence of 
catalyst particles comprised of one or more Group VIII (IU 8-10) noble 
metals (e.g., platinum, iridium, rhodium, palladium) and a halogen 
combined with a porous carrier, such as a refractory inorganic oxide. 
In a common form, the reforming process will employ the catalyst particles 
in several reaction zones interconnected in a series flow arrangement. 
There may be any number of reaction zones, but usually the number of 
reaction zones is 3, 4 or 5. Because reforming reactions occur generally 
at an elevated temperature, and because reforming reactions generally are 
endothermic, each reaction zone usually has associated with it one or more 
heating zones, which heat the reactants to the desired reaction 
temperature and which supply the endothermic heat of reaction for the 
reaction zones. As a consequence of these considerations, the most common 
process flow through the train of heating and reaction zones in a 
3-reactor catalytic reforming processes is as follows. 
A naphtha-containing feedstock admixes with a hydrogen-containing recycle 
gas to form a combined feed stream, which passes through a combined feed 
heat exchanger. In the combined feed heat exchanger, the combined feed is 
heated by exchanging heat with the effluent of the third reactor. The 
heating of the combined feed stream that occurs in the combined feed heat 
exchanger is, however, insufficient to heat the combined feed stream to 
the desired inlet temperature of the first reactor. Consequently, after 
exiting the combined feed heat exchanger and prior to entering the first 
reactor, the combined feed stream requires additional heating. This 
additional heating occurs in a heater, which is commonly referred to as a 
charge heater, which heats the combined feed stream to the desired inlet 
temperature of the first reactor. 
The combined feed stream then passes to and through the first reactor. 
Because of the endothermic reforming reactions that occur in the first 
reactor, the temperature of the effluent of the first reactor falls not 
only to less than the temperature of the combined feed to the first 
reactor, but also and more importantly, to less than the desired inlet 
temperature of the second reactor. Therefore, the effluent of the first 
reactor passes through another heater, which is commonly referred to as 
the first interheater and which heats the first reactor effluent to the 
desired inlet temperature of the second reactor. 
On exiting the first interheater the first reactor effluent enters the 
second reactor. As in the first reactor, endothermic reactions cause 
another decline in temperature across the second reactor. Generally, 
however, the temperature decline across the second reactor is less than 
the temperature decline across the first reactor, because the reactions 
that occur in the second reactor are generally less endothermic than the 
reactions that occur in the first reactor. Despite the somewhat lower 
temperature decline across the second reactor, the effluent of the second 
reactor is nevertheless still at a temperature that is less than the 
desired inlet temperature of the third reactor. Consequently, the effluent 
of the second reactor passes through another heater, which is commonly 
referred to as the second interheater, and then passes to the third 
reactor. 
In the third reactor, endothermic reactions cause yet another temperature 
decline, which is generally less than that across the second reactor, for 
the like reason that the temperature decline across the second reactor is 
generally less than that across the first reactor. The effluent of the 
third reactor passes to the previously mentioned combined feed exchanger, 
where the effluent of the third reactor is cooled by exchanging heat with 
the combined feed stream. 
After a period of time in use, the catalyst becomes deactivated during the 
course of reforming reactions as a result of mechanisms such as the 
deposition of coke on the particles. Coke is comprised primarily of 
carbon, but is also comprised of a small quantity of hydrogen. Coke 
decreases the ability of catalyst to promote reforming reactions to the 
point that continued use of the catalyst is no longer practical or 
economical. At that point, the catalyst must be reconditioned, or 
regenerated, before it can be reused in a reforming process. 
Numerous regeneration methods are in use commercially, and nearly all 
involve to some extent the combustion of coke from the surface of the 
catalyst. The particular method of regeneration that a specific reforming 
process employs depends on the design of the catalyst bed(s) in the 
reforming reactor(s). A commercial reforming reactor generally employs one 
of two different designs of catalyst beds: moving beds and fixed beds. In 
a moving bed, deactivated catalyst is withdrawn from the catalyst bed, 
while fresh or regenerated catalyst is added to the bed. Moving catalyst 
beds allow catalyst to be continuously moved from the reactor to an 
adjacent regeneration zone, regenerated, and moved back to the reactor. 
This is commonly referred to as continuous regeneration, although in 
practice it is often semicontinuous. 
In contrast, fixed catalyst beds keep the catalyst stationary. When the 
catalyst in a fixed bed reactor becomes deactivated, the reactor is 
temporarily taken out of service while the catalyst is either regenerated 
in situ or else unloaded and replaced with regenerated or fresh catalyst. 
Two types of fixed bed regeneration methods are used commercially: cyclic 
regeneration and semi-regeneration. In the cyclic regeneration method, at 
least one or at most not all of the reactors are taken out of service at 
any one time and the reforming process continues in operation with the 
remaining reactors. After the deactivated catalyst is regenerated, the 
reactor is placed back in service, which in turn allows another reactor to 
be taken out of service for regeneration. In semi-regenerative reforming, 
the reforming process is temporarily stopped and all of the reactors are 
taken out of service simultaneously for regeneration. After the catalyst 
has been regenerated, all the reactors are placed back in service and the 
reforming process is resumed. 
Commercial reforming process units that use the flow schemes and 
regeneration methods just described require a large investment of capital, 
and one of the major capital costs is that associated with the charge 
heater. As previously mentioned, the charge heater heats the combined feed 
stream to the desired inlet temperature of the first reactor. As also 
mentioned, the charge heater is typically needed because the heat that is 
transferred by the combined feed exchanger from the third reactor effluent 
to the combined feed is not sufficient to heat the combined feed stream to 
the desired inlet temperature of the first reactor. But, because of the 
capital cost of the charge heater, reforming processes are sought that 
eliminate the need for a charge heater. 
It is known that a reforming process unit that uses the semi-regeneration 
method can operate without a charge heater. In one such unit, the combined 
feed heat exchanger comprises a plate type heat exchanger which heats the 
combined feed stream, and the heated combined feed stream then passes 
directly to a first reactor which contains a fixed bed of platinum-rhenium 
catalyst. The first reactor effluent then passes to a train of pairs of 
heaters and reactors that contain fixed beds of platinum-rhenium catalyst; 
that is to a first heater, a second reactor, a second heater, a third 
reactor, a third heater, and a fourth reactor. The fourth reactor effluent 
transfers heat to the combined feed stream via the plate type heat 
exchanger, and then passes to a product recovery section. The inlet 
temperature of the first reactor ranges typically from about 784 to about 
849.degree. F. (418 to 454.degree. C.). The inlet temperatures of the 
second, third, and fourth reactors typically are within about 5.degree. F. 
(3.degree. C.) of each other, and are typically from about 81 to about 
102.degree. F. (45 to 57.degree. C.) hotter than the inlet temperature of 
the first reactor, and usually vary over the range of from about 878 to 
about 939.degree. F. (470 to 504.degree. C.). The molar ratio of hydrogen 
per hydrocarbon feedstock is typically between 4 and 6. 
It is also known that a reforming unit that uses the continuous 
regeneration method can operate with different feed inlet temperatures for 
each of the reactors. Typically, such a unit has a train of three, four or 
five pairs of heaters and reactors that contain moving beds of catalyst, 
but many of the various possible combinations of different inlet 
temperatures, which together form what is usually called the temperature 
profile of the unit, are perhaps best illustrated with a three-reactor 
unit. If the inlet temperatures of all three reactors are the same, then 
the temperature profile is commonly called flat, which is the profile that 
is most frequently employed in reforming units using continuous 
regeneration. If the inlet temperature of the first reactor is less than 
the inlet temperature of the second reactor, which is in turn less than 
the inlet temperature of the third reactor, then the profile of the 
reactor inlet temperatures is usually said to be ascending. If the first 
inlet temperature is more than the second inlet temperature, which is more 
than the third inlet temperature, then the profile is normally called 
descending. If the second inlet temperature is more than both the first 
and third inlet temperatures, then the profile is often said to resemble a 
hill. If the second inlet temperature is less than both the first and 
third inlet temperatures, then the profile is frequently said to look like 
a valley. Thus, in the ascending and hill profiles, the inlet temperature 
of the second reactor is greater than the inlet temperature of the third 
reactor. 
The most common reason for operating with a non-flat (i.e., skewed) reactor 
inlet temperature profile is to allocate the required heat duty among the 
heaters in the heater-reactor train. Ideally, all of the heaters are 
individually delivering heat at approximately the same percentage of their 
individual design duties. When each heater is operating at the same 
percentage of its design duty as any other heater in the train is 
operating as a percentage of that other heater's design duty, then the 
heater duties are said to be "balanced." Of course, a heater should not, 
as a general rule, be operated in excess of its design duty, that is the 
percentage should generally be less than or equal to 100%. A flat profile 
could result in imbalance of the operating duties of the heaters in the 
train, if some of the operating variables such as feedstock quality or 
throughput differ significantly from their design values, or if flow 
maldistribution or mechanical problems causes the performance of a reactor 
to fall significantly below its expected performance. 
An illustration of attempting to balance heater duties in a commercial 
continuous reforming process by skewing reactor inlet temperatures is 
described in the article by Richard Lee, et al. entitled "Reforming 
Processes, Maximizing Profitability," which begins at page 151 in Volume 
47 of the Encyclopedia of Chemical Processing and Design, edited by John 
J. McKetta and published by Marcel Dekker, Inc., New York in 1994. In the 
example in the article by Lee et al., a valley-shaped profile of reactor 
inlet temperatures is recommended, where the inlet temperatures of 
parallel reactors 1 and 2 are the same and greater than the inlet 
temperature of reactor 3, which is less than the inlet temperature of 
reactor 4. Reactor 4's inlet temperature may be the same as or less than 
that of reactors 1 and 2. The largest difference between the reactor 1, 2, 
and 4 inlet temperatures and the reactor 3 inlet temperature is 26.degree. 
F. (14.degree. C.). The Lee et al. article also teaches that the magnitude 
of the differences between the gasoline range product (C.sub.5 + yield) 
when running an equal (that is, flat) reactor inlet temperature profile 
versus a staggered (that is, skewed) reactor inlet temperature profile is 
expected to be no more than 0.5% of feed. 
Operating a reforming process without heating between the combined feed 
heat exchanger and the first or lead reactor has certain benefits which, 
however, are incapable of being realized in a reforming process unit that 
uses the semi-regeneration method. Therefore, reforming processes are 
sought that are capable of achieving more fully all the possible benefits 
and advantages of operating without a charge heater. 
SUMMARY OF THE INVENTION 
This invention is a reforming process that employs moving catalyst particle 
beds for the first, or lead, reactor and at least one other reactor, 
wherein the other reactor receives hydrocarbon effluent and catalyst 
particles from the first reactor and has an inlet temperature that is at 
least 60.degree. F. (33.degree. C.) hotter than the inlet temperature of 
the first reactor. Even though the inlet temperature of the first reactor 
is low in comparison to that of the other reactor, reforming reactions 
nevertheless occur in the first reactor. The reactions that take place in 
the first reactor reform mainly naphthenes, such as the dehydrogenation of 
cyclohexanes and also the dehydroisomerization of alkylcyclopentanes. 
These reactions occur readily at reforming conditions, and both reactions 
yield aromatics as well as by-product hydrogen. Moreover, these reactions 
occur at a relatively low temperature, which inhibits coke formation in 
the first reactor. Therefore, the catalyst particles that leave the first 
reactor of this invention have a relatively low coke content in comparison 
to prior art processes where the first reactor operates at a relatively 
higher temperature. 
In this invention, the low-coke-containing catalyst particles from the 
first reactor enter the other reactor, which as already mentioned, 
operates at a relatively high inlet temperature relative to the first 
reactor. In the other reactor, reforming reactions occur, such as 
dehydrocyclization of paraffins and of olefins to yield aromatics, which 
are comparatively more difficult than the naphthene reforming reactions 
that occur in the first reactor and require relatively hotter 
temperatures. These comparatively difficult reactions, and the hotter 
temperatures, tend to accelerate coke formation. Thus, in this invention 
these comparatively difficult reactions occur in the presence of catalyst 
particles having a relatively low coke content, whereas in the prior art 
these reactions take place in the presence of catalyst particles having a 
relatively high coke content. By-product hydrogen from the first reactor 
also helps to inhibit coke formation in the other reactor in this 
invention. 
Accordingly, an advantage of this invention over the prior art 
semi-regenerative process is a reduction in the coke content on the 
catalyst particles leaving the first reactor and entering the other 
reactor. This advantage is more pronounced at lower hydrogen recycle 
ratios to the first reactor. This advantage may in some cases in turn 
decrease the coke content on the catalyst leaving the second reactor, 
which can be easily exploited to increase the profitability of a reforming 
process. For example, if the regeneration section of a reforming process 
is already operating at its maximum coke-burning capacity, this invention 
may allow the reactor section, in particular the other reactor, to be 
operated at conditions that would not have been considered practical 
because of their effect on coke formation. Thus with this invention, 
feedstock costs could be decreased by charging a less-expensive paraffinic 
naphtha to the reforming process, a more valuable product could be 
produced by increasing reformate octane, or utilities costs could be 
reduced by decreasing the recycle gas ratio. By operating at these new 
conditions, the reforming process would become more profitable. 
Alternatively, even if no changes were made in the operating conditions of 
the reforming process, this invention may allow the frequency of catalyst 
regeneration to be decreased. Less frequent regenerations would slow the 
deterioration of the select catalytic properties (surface area, metal 
dispersion, etc.) that occur as a result of the frequent regenerations. 
This would in turn prolong the period of time between catalyst changeouts 
and decrease the catalyst replacement costs. For catalysts that have 
suffered deterioration in surface area, this invention would help to 
maintain high catalyst activity by decreasing the average coke content of 
the catalyst in the reactors for the same operating conditions. 
The temperature difference between the first reactor and the other reactor 
is most easily and most advantageously achieved by passing the combined 
feed stream from the combined feed exchanger to the first reactor without 
any intermediate heating. Thus, in one embodiment this invention is a 
reforming process in which the combined feed stream from the combined feed 
exchanger passes, without any additional heating, to the first reactor. 
This embodiment is, therefore, a reforming process that does not require a 
charge heater. Accordingly, the heat that is recovered from the last 
reactor effluent and transferred to the combined feed by the combined feed 
exchanger is sufficient to heat the combined feed stream to the desired 
inlet temperature of the first reactor. 
One way to think about this embodiment of the invention is in comparison 
with a prior art reforming process that uses n reactors. From this point 
of view, the principal benefit of this embodiment is the reduction in 
capital cost by eliminating the charge heater. That is, this embodiment 
requires no more than n-1 heaters, rather than the n heaters used in the 
prior art process. The savings in capital cost include the cost not only 
of the mechanical components of the charge heater itself, such as the 
heater tubes, refractory linings, outer shell, and foundation, but also of 
the associated electrical instrumentation, and of the interconnecting 
piping both upstream from the combined feed exchanger and downstream to 
the first reactor. In addition, in some embodiments of this invention 
elimination of the charge heater reduces the pressure drop of the reactor 
section, which decreases the energy-related operating cost of the recycle 
compressor, as well as the capital cost of the recycle compressor and its 
associated motor or turbine. Also, elimination of the charge heater 
decreases the plot space that would have been occupied by the charge 
heater, thereby freeing that plot space for other more profitable or 
productive uses. 
Another way of looking at this embodiment of the invention is in comparison 
with a reforming process that uses n heaters. In this light, another 
benefit of eliminating the charge heater is an increase in the efficient 
use of the catalyst, by employing n+1 reactors, as opposed to only n 
reactors in the prior art processes. As a general rule, given the same 
amount of catalyst in a reforming process, as the number of reactors 
increases, the efficiency of use of the catalyst also increases. So, for 
example, a 5-reactor reforming process uses a given amount of catalyst 
more efficiently than a 4-reactor reforming process, which in turn uses 
that same given amount of catalyst more efficiently than a 3-reactor 
system. Thus, for a given amount of catalyst and a given number of 
heaters, this embodiment employs the catalyst more efficiently than the 
prior art processes. 
Yet another way of thinking about this embodiment of the invention is in 
comparison with two prior art processes, namely a prior art process that 
uses n reactors and n heaters and another prior art process that uses n+1 
reactors and n+1 heaters. For a given quality and quantity of feedstock 
and a given reformate product quality, a refiner who wished to build a new 
moving bed reforming process unit was forced to choose between only two 
options: n reactors and n heaters, or n+1 reactors and n+1 heaters. For 
most refiners, the value of n is 3, and so the refiner had to choose 
between 3 reactors and 3 heaters, or 4 reactors and 4 heaters. The refiner 
made the final decision based on an economic evaluation of the total 
return on his investment both in catalyst on the one hand and in equipment 
and machinery on the other. 
Unfortunately, it has often been the case that neither a 3 reactor/3 heater 
process nor a 4 reactor/4 heater process is the economic optimum, and the 
refiner has had to choose the less-worse of two non-optimum choices. The 
advantage of the embodiment of this invention that eliminates the charge 
heater, then, is to give the refiner a new, third choice, such as a 4 
reactor/3 heater process, that is intermediate between the two prior art 
choices and is possibly closer to the refiner's economically optimum 
process unit. This invention does not necessarily allow the refiner to 
avoid undertaking an economic evaluation such as that used to evaluate the 
prior art processes. Instead, this invention gives the refiner an 
additional option for the outcome of that evaluation that may in some 
cases be the most profitable option of all. 
Accordingly, this invention is in its broadest terms is a process for 
reforming hydrocarbons. A feedstock comprising hydrocarbons is passed to a 
first catalyst bed that contains catalyst particles. The feedstock enters 
the first catalyst bed at a first inlet temperature. In the first catalyst 
bed, hydrocarbons contact the catalyst particles and are reformed. A first 
bed effluent stream comprising hydrocarbons is withdrawn from the first 
catalyst bed. At least a portion of the first bed effluent stream passes 
to a second catalyst bed which contains catalyst particles. At least a 
portion of the first bed effluent stream enters the second catalyst bed at 
a second inlet temperature that is at least about 60.degree. F. greater 
than the first inlet temperature. In the second catalyst bed, hydrocarbons 
are contacted with catalyst particles and reformed. A second bed effluent 
stream comprising reformate is withdrawn from the second catalyst bed. 
Reformate is recovered from the second bed effluent stream. At least 
periodically, catalyst particles move through the first catalyst bed and 
the second catalyst bed by withdrawing catalyst particles from the second 
catalyst bed, passing catalyst particles from the first catalyst bed to 
the second catalyst bed, and adding catalyst particles to the first 
catalyst bed. 
In another embodiment, this invention is a process for reforming 
hydrocarbons. A combined feed stream comprising hydrocarbons and hydrogen 
passes to a combined feed heat exchanger, which heats the combined feed 
stream to produce a heated feed stream that comprises hydrocarbons and 
hydrogen. Without heating, at least a portion of the heated feed stream 
passes to a first reactor at a combined feed temperature. In the first 
reactor, hydrocarbons contact catalyst particles in a first catalyst bed 
and are reformed. A first reactor effluent stream comprising hydrocarbons 
is withdrawn from the first reactor. At least a portion of the first 
reactor effluent stream passes to a first heater, which heats the stream 
to produce a second reactor feed stream comprising hydrocarbons. At least 
a portion of the second reactor feed stream passes to a second reactor at 
a second reactor feed temperature that is at least about 60.degree. F. 
greater than the combined feed temperature. In a second catalyst bed in 
the second reactor, hydrocarbons contact catalyst particles and are 
reformed. A second reactor effluent stream comprising hydrocarbons is 
withdrawn from the second reactor. At least a portion of the second 
reactor effluent stream passes to the combined feed heat exchanger. In the 
combined feed heat exchanger, heat is exchanged from the portion of the 
second reactor effluent stream to the combined feed stream. A combined 
feed exchanger effluent stream comprising reformate is withdrawn from the 
combined feed heat exchanger. Reformate is recovered from the combined 
feed exchanger effluent stream. Catalyst particles are at least 
periodically moved through the first catalyst bed and the second catalyst 
bed by withdrawing catalyst particles from the second catalyst bed, 
passing catalyst particles from the first catalyst bed to the second 
catalyst bed, and adding catalyst particles to the first catalyst bed. 
INFORMATION DISCLOSURE 
Catalytic reforming is described in Chapter 4.1, "UOP Platforming Process," 
of the book entitled Handbook of Petroleum Refining Processes, Second 
Edition, edited by Robert A. Meyers, published by McGraw-Hill Book 
Company, in New York, in 1997. The teachings of Chapter 4.1 of Meyers' 
book are incorporated herein by reference. 
U.S. Pat. No. 3,652,231 (Greenwood et al.) describes a process for the 
catalytic reforming of hydrocarbons with continuous catalyst regeneration 
that employs an apparatus that uses moving catalyst beds. U.S. Pat. No. 
3,647,680 (Greenwood et al.) and U.S. Pat. No. 3,692,496 (Greenwood et 
al.) also describe the regeneration of reforming catalyst. The teachings 
of U.S. Pat. Nos. 3,652,231, 3,647,680, and 3,692,496 are incorporated 
herein by reference. 
U.S. Pat. No. 4,325,806 (Peters) discloses a hydrocarbon conversion process 
having at least three reaction zones, wherein one portion of the effluent 
of the first reaction zone is passed to the second reaction zone, and 
another portion of the effluent of the first reaction zone is combined 
with the entire effluent of the second reaction zone and passed to a third 
reaction zone. 
U.S. Pat. No. 4,325,807 (Peters) discloses a hydrocarbon conversion process 
having at least four reaction zones, where the first and second reaction 
zones are in a parallel-flow arrangement. In addition, the combined 
effluent of the first and second reaction zones is divided into two 
portions: one portion is passed to the third reaction zone, and the other 
portion is combined with the entire effluent of the third reaction zone 
and passed to the fourth reaction zone.

DETAILED DESCRIPTION OF THE INVENTION 
The hydrocarbon feedstock that is charged to this invention will comprise 
naphthenes and paraffins that boil within the gasoline range. The 
preferred charge stocks are naphthas consisting principally of naphthenes 
and paraffins, although, in many cases, aromatics also will be present. 
This preferred class includes straight-run gasolines, natural gasolines, 
synthetic gasolines, and the like. As an alternative embodiment, it is 
frequently advantageous to charge thermally or catalytically cracked 
gasolines or partially reformed naphthas. Mixtures of straight-run and 
cracked gasoline-range naphthas can also be used to advantage. The 
gasoline-range naphtha charge stock may be a full-boiling gasoline having 
an initial boiling initial boiling point of from about 104 to about 
180.degree. F. (40 to 82.degree. C.) and an end boiling point within the 
range of from about 320 to about 428.degree. F. (160 to 220.degree. C.), 
or may be a selected fraction thereof which generally will be a 
higher-boiling fraction commonly referred to as a heavy naphtha--for 
example, a naphtha boiling in the range of from about 212 to about 
392.degree. F. (100 to 200.degree. C.). In some cases, it is also 
advantageous to charge pure hydrocarbons or mixtures of hydrocarbons that 
have been recovered from extraction units--for example, raffinates from 
aromatics extraction or straight-chain paraffins--which are to be 
converted to aromatics. In some other cases, the feedstock may also 
contain light hydrocarbons that have from 1 to 5 carbon atoms, but since 
these light hydrocarbons cannot be readily reformed into aromatic 
hydrocarbons, these light hydrocarbons entering with the feedstock are 
generally minimized. 
This invention is applicable to the catalytic reforming of hydrocarbons in 
a reforming reaction system having at least two catalytic reaction zones 
where at least a portion of the reactant stream and at least a portion of 
the catalyst particles flow serially through the reaction zones. Reaction 
systems having multiple zones generally take one of two forms, a 
side-by-side form or a stacked form. In the side-by-side form, multiple 
and separate reaction vessels, each comprising a reaction zone, are placed 
along side each other. In the stacked form, one common reaction vessel 
contains the multiple and separate reaction zones that are placed on top 
of each other. 
Although the reaction zones can comprise any number of arrangements for 
hydrocarbon flow such as downflow, upflow, and crossflow, the most common 
reaction zones to which this invention is applied are radial flow. A 
radial flow reaction zone generally consists of cylindrical sections 
having varying nominal cross-sectional areas, vertically and coaxially 
disposed to form the reaction zone. Briefly, a radial flow reaction zone 
typically comprises a cylindrical reaction vessel containing a cylindrical 
outer catalyst retaining screen and a cylindrical inner catalyst retaining 
screen that are both coaxially-disposed with the reaction vessel. The 
inner screen has a nominal, internal cross-sectional area that is less 
than that of the outer screen, which has a nominal, internal 
cross-sectional area that is less than that of the reaction vessel. The 
reactant stream is introduced into the annular space between the inside 
wall of the reaction vessel and the outside surface of the outer screen. 
The reactant stream passes through the outer screen, flows radially 
through the annular space between the outer screen and the inner screen, 
and passes through the inner screen. The stream that is collected within 
the cylindrical space inside the inner screen is withdrawn from the 
reaction vessel. Although the reaction vessel, the outer screen, and the 
inner screen may be cylindrical, they may also take any suitable shape, 
such as triangular, square, oblong, and diamond, depending on many design, 
fabrication, and technical considerations. For example, it is common for 
the outer screen to not be a continuous cylindrical screen but to instead 
be an arrangement of separate, elliptical, tubular screens called scallops 
that are arrayed around the circumference of the inside wall of the 
reaction vessel. The inner screen is commonly a perforated centerpipe that 
is covered around its outer circumference with a screen. 
In this invention, the reformer employs at least two moving catalyst beds 
for reforming hydrocarbons. It is an essential aspect of this invention 
that catalyst particles withdrawn from the first moving catalyst bed pass 
to the second moving catalyst bed. Preferably, each moving bed is a moving 
packed bed and all of the catalyst particles withdrawn from the first bed 
pass to second bed. This invention is applicable to catalytic reforming 
processes wherein the catalyst comprises particles that are movable 
through at least two reaction zones. The catalyst particles are movable 
through the reaction zones by any of a number of motive devices including 
conveyors or transport fluid, but most commonly the catalyst particles are 
movable through the reaction zone by the force of gravity. Typically, in a 
radial flow reaction zone the catalyst particles fill the annular space 
between the inner and outer screens, which is called the catalyst bed. 
Catalyst particles are withdrawn from a bottom portion of a reaction zone, 
and catalyst particles are introduced into a top portion of the reaction 
zone. The catalyst particles withdrawn from a reaction zone can 
subsequently be recovered from the process, regenerated in a regeneration 
zone of the process, or transferred to another reaction zone. Likewise, 
the catalyst particles added to a reaction zone can be catalyst that is 
being newly added to the process, catalyst that has been regenerated in a 
regeneration zone within the process, or catalyst that is transferred from 
another reaction zone. 
Illustrative reaction vessels that have stacked reaction zones and that may 
be used to practice this invention are shown in U.S. Pat. Nos. 3,706,536 
(Greenwood, et al.) and 5,130,106 (Greenwood, et al.), the teachings of 
which are incorporated herein by reference. Transfer of the 
gravity-flowing catalyst particles from one reaction to another, the 
introduction of fresh or regenerated catalyst particles, and the 
withdrawal of coke-containing spent catalyst particles is effected through 
catalyst transfer conduits. 
The reforming reactions that occur in the present invention are normally 
effected in the presence of catalyst particles comprised of one or more 
Group VIII (IU 8-10) noble metals (e.g., platinum, iridium, rhodium, 
and palladium) and a halogen combined with a porous carrier, such as a 
refractory inorganic oxide. U.S. Pat. No. 2,479,110, for example, teaches 
an alumina-platinum-halogen reforming catalyst. Although the catalyst may 
contain 0.05-2.0 wt-% of Group VIII metal, this invention is most 
economical when practiced in the presence of a less expensive catalyst, 
such as a catalyst containing 0.05-0.5 wt-% of Group VIII metal. The 
preferred noble metal is platinum. In addition, the catalyst may contain 
indium and/or a lanthanide series metal such as cerium. The catalyst 
particles may also contain 0.05-0.5 wt-% of one or more Group IVA (IU 
14) metals (e.g., tin, germanium, and lead), such as described in U.S. 
Pat. No. 4,929,333, U.S. Pat. No. 5,128,300, and the references cited 
therein. The halogen is normally chlorine. Alumina is a commonly used 
carrier. The preferred alumina materials are known as gamma, eta, and 
theta alumina, with gamma and eta alumina giving the best results. An 
important property related to the performance of the catalyst is the 
surface area of the carrier. Preferably, the carrier will have a surface 
area of from about 100 to about 500 m.sup.2 /g. The activity of catalysts 
having a surface area of less than about 130 m.sup.2 /g is known to be 
more detrimentally affected by catalyst coke than catalysts having a 
higher surface area. Accordingly, the benefit of this invention is 
believed to be greater for processes that use a catalyst having a surface 
area of less than about 130 m.sup.2 /g, because of the reduction in coke 
content of the catalyst leaving the first reactor. The particles are 
usually spheroidal and have a diameter of from about 1/16.sup.th to about 
1/8.sup.th inch (1.6 to 3.1 mm), though they may be as large as 1/4.sup.th 
inch (6.35 mm) or as small as 1/24.sup.th inch (1.06 mm). In a particular 
reforming reactor, however, it is desirable to use catalyst particles 
which fall in a relatively narrow size range. A preferred catalyst 
particle diameter is 1/16.sup.th inch (1.6 mm). 
In preferred form, the reforming process will employ a moving bed reaction 
vessel and a moving bed regeneration vessel, and the present invention is 
applicable to such a reforming process. Regenerated catalyst particles are 
fed to the reaction vessel, which is typically comprised of several 
reaction zones, and the particles flow through the reaction vessel by 
gravity. Catalyst is withdrawn from the bottom of the reaction vessel and 
transported to the regeneration vessel. In the regeneration vessel, a 
multi-step regeneration process is typically used to regenerate the 
catalyst to restore its full ability to promote reforming reactions. U.S. 
Pat. Nos. 3,652,231 (Greenwood et al.), 3,647,680 (Greenwood et al.) and 
3,692,496 (Greenwood et al.) describe catalyst regeneration vessels that 
are suitable for use in a reforming process. Catalyst flows by gravity 
through the various regeneration steps and then is withdrawn from the 
regeneration vessel and transported to the reaction vessel. Arrangements 
are provided for adding fresh catalyst as make-up to and for withdrawing 
spent catalyst from the process. Movement of catalyst through the reaction 
and regeneration vessels is often referred to as continuous though, in 
practice, it is semicontinuous. By semicontinuous movement it is meant the 
repeated transfer of relatively small amounts of catalyst at closely 
spaced points in time. For example, one batch every twenty minutes may be 
withdrawn from the bottom of the reaction vessel and withdrawal may take 
five minutes, that is, catalyst will flow for five minutes. If the 
catalyst inventory in a vessel is relatively large in comparison with this 
batch size, the catalyst bed in the vessel may be considered to be 
continuously moving. A moving bed system has the advantage of maintaining 
production while the catalyst is removed or replaced. 
As mentioned previously, this invention is a new moving bed reforming 
process that may be a more economical or profitable option than any of the 
prior art moving bed reforming processes. Whether this invention is the 
most economical or profitable option for a given refiner requires an 
economic evaluation of the refiner's total return on investment in 
catalyst and equipment. A complete economic evaluation is not necessary to 
practice this invention and therefore does not need to be described in 
detail herein. However, such an economic evaluation will generally take 
into account a large number of factors that are interrelated in a complex 
manner. These factors include feedstock composition (e.g., paraffins, 
naphthenes, and aromatics in the feedstock), liquid hourly space velocity, 
recycle gas ratio, reactor inlet temperatures, reactor outlet 
temperatures, catalyst performance, catalyst cost, reactor fabrication 
cost, minimum and maximum reactor dimensions, from the viewpoint of ease 
in fabrication as well as of ease in maintenance, and reformate octane. 
The reaction zones of the present invention are operated at reforming 
conditions, which include a range of pressures generally from atmospheric 
pressure (0 psi(g)) to 1000 psi(g) (0 to 6895 kPa(g)), with particularly 
good results obtained at the relatively low pressure range of from about 
40 to about 200 psi(g) (276 to 1379 kPa(g)). The overall liquid hourly 
space velocity (LHSV) based on the total catalyst volume in all of the 
reaction zones is generally from 0.1 to 10 hr.sup.-1, preferably from 
about 1 to about 5 hr.sup.-1, and more preferably from about 1.5 to about 
2.0 hr.sup.-1. A relatively low LHSV is preferred because a process 
operating at a low LHSV employs a relatively large volume of catalyst per 
volume of feedstock that is charged to the reforming process, and in 
general the greater the volume of catalyst that is present in a process, 
the more readily it is that a portion of the catalyst can operate at a 
relatively low temperature while the process nevertheless achieves the 
desired reformate product quality. 
Generally, hydrogen is supplied to provide an amount of from about 1 to 
about 20 moles of hydrogen per mole of hydrocarbon feedstock entering the 
reforming zone. Hydrogen is preferably supplied to provide an amount of 
less than about 3.5 moles of hydrogen per mole of hydrocarbon feedstock 
entering the reforming zone. If hydrogen is supplied, it may be supplied 
upstream of the combined feed exchanger, downstream of the combined feed 
exchanger, or both upstream and downstream of the combined feed exchanger. 
Providing hydrogen with the hydrocarbon feedstock is not, however, a 
requirement of this invention, and thus it is possible that no hydrogen 
may be supplied to enter the reforming zone with the hydrocarbon 
feedstock. Even if hydrogen is not provided with the hydrocarbon feedstock 
to the first reaction zone, the naphthene reforming reactions that occur 
within the first reaction zone yield hydrogen as a by-product. This 
by-product, or in-situ-produced, hydrogen leaves the first reaction zone 
in an admixture with the first reaction zone effluent and then becomes 
available as hydrogen to the second reaction zone and other downstream 
reaction zones. This in situ hydrogen in the first reaction zone effluent 
usually amounts to between 0.5 and 2 moles of hydrogen per mole of 
hydrocarbon feedstock. 
It is believed that one of the benefits of this invention, namely a 
reduction in the formation of coke on the catalyst, is more pronounced as 
the molar ratio of hydrogen per hydrocarbon feedstock to the first reactor 
decreases. Thus, it is expected that, as the molar ratio decreases from 
3.5 to say 2.0, or to even lower ratios such as 1.0 or 0.5, a decrease in 
the inlet temperature of the first reactor significantly decreases the 
coke formation in the first reactor relative to the first reactor of the 
prior processes that operate at a higher temperature. Furthermore, it is 
believed that any increased coke formation that occurs in the second, 
third, and/or subsequent reactors of this invention will be more than 
offset by the decreased coke formation in the first reactor, so that the 
overall coke make in this invention will be less than that in the prior 
art processes. In addition, it should be pointed out that the lower coke 
contents of the catalyst that passes through the second, third, and/or 
subsequent reactors in this invention can have a beneficial affect on the 
performance of some catalysts, such as catalysts having low surface area, 
in those reactors, particularly from the point of view of catalyst 
activity. 
Typically, the rate of catalyst movement through the catalyst beds may 
range from as little as 200 pounds (90.7 kg) per hour to 4500 pounds (2041 
kg) per hour, or more. 
The combined feed stream, or the hydrocarbon feedstock if no hydrogen is 
provided with the hydrocarbon feedstock, enters a heat exchanger at a 
temperature of generally from about 150 to about 350.degree. F. (65 to 
177.degree. C.), and more usually from about 200 to about 250.degree. F. 
(93 to 121.degree. C.). Because hydrogen is usually provided with the 
hydrocarbon feedstock, this heat exchanger may be referred to herein as 
the combined feed heat exchanger, even if no hydrogen is supplied with the 
hydrocarbon feedstock. The combined feed heat exchanger heats the combined 
feed stream by transferring heat from the effluent stream of the last 
reforming reactor to the combined feed stream. The combined feed heat 
exchanger is preferably an indirect, rather than a direct, heat exchanger, 
in order to prevent valuable reformate product in the last reactor's 
effluent from intermixing with the combined feed and thereby being 
recycled to the reforming reactors, where the reformate quality could be 
degraded. 
Although the flow pattern of the combined feed stream and the last reactor 
effluent stream within the combined feed heat exchanger could be 
completely-cocurrent, reversed, mixed, or cross flow, the flow pattern is 
preferably countercurrent. By a countercurrent flow pattern, it is meant 
that the combined feed stream, while at its coldest temperature, contacts 
one end (i.e., the cold end) of the heat exchange surface of the combined 
feed heat exchanger while the last reactor effluent stream contacts the 
cold end of the heat exchange surface at its coldest temperature also. 
Thus, the last reactor effluent stream, while at its coldest temperature 
within the heat exchanger, exchanges heat with the combined feed stream 
that is also at its coldest temperature within the heat exchanger. At 
another end (i.e., the hot end) of the combined feed heat exchanger 
surface, the last reactor effluent stream and the combined feed stream, 
both at their hottest temperatures within the heat exchanger, contact the 
hot end of the heat exchange surface and thereby exchange heat. Between 
the cold and hot ends of the heat exchange surface, the last reactor 
effluent stream and the combined feed stream flow in generally opposite 
directions, so that, in general, at any point along the heat transfer 
surface, the hotter the temperature of the last reactor effluent stream, 
the hotter is the temperature of combined feed stream with which the last 
reactor effluent stream exchanges heat. For further information on flow 
patterns in heat exchangers, see, for example, pages 10-24 to 10-31 in 
Perry's Chemical Engineers' Handbook, Sixth Edition, edited by Robert H. 
Perry et al., published by McGraw-Hill Book Company in New York, in 1984, 
and the references cited therein. 
In this invention, the combined feed heat exchanger generally operates with 
a hot end approach that is generally less than about 100.degree. F. 
(56.degree. C.), and preferably less than about 60.degree. F. (33.degree. 
C.), and more preferably less than about 50.degree. F. (28.degree. C.). As 
used herein, the term "hot end approach" is defined as follows, based on a 
heat exchanger that exchanges heat between a hotter last reactor effluent 
stream and a colder combined feed stream, where T1 is the inlet 
temperature of the last reactor effluent stream, T2 is the outlet 
temperature of the last reactor effluent stream, t1 is the inlet 
temperature of the combined feed stream, and t2 is the outlet temperature 
of the combined feed stream. Then, as used herein, for a countercurrent 
heat exchanger the "hot end approach" is defined as the difference between 
T1 and t2. In general, the smaller the hot end approach, the greater is 
the degree to which the heat in the last reactor's effluent is exchanged 
to the combined feed stream, and the lesser is the need for a charge 
heater. 
Although some shell-and-tube type heat exchangers are capable of achieving 
the hot end approaches set forth in the prior two paragraphs, the 
preferred type of heat exchanger for use in this invention is a plate type 
heat exchanger. Plate type exchangers are well known and commercially 
available in several different and distinct forms, such as spiral, plate 
and frame, brazed-plate fin, and plate fin-and-tube types. Plate type 
exchangers are described generally at pages 11-21 to 11-23 in Perry's 
Chemical Engineers' Handbook, Sixth Edition, edited by R. H. Perry et al., 
and published by McGraw Hill Book Company, in New York, in 1984. 
The combined feed stream leaves the combined feed heat exchanger at a 
temperature of generally from about 750 to about 960.degree. F. (399 to 
516.degree. C.). Because in one embodiment of this invention there is no 
heating between the combined feed heat exchanger and the lead or first 
reactor, the outlet temperature of the combined feed stream from the 
combined feed heat exchanger is essentially the same as the temperature of 
the inlet temperature of the first reforming reactor. However, as a 
practical matter, there may be some loss of heat between the outlet of the 
combined feed exchanger and the inlet of the first reactor, and therefore 
in this context, the expression "essentially the same" means that the 
inlet temperature of the first reforming reactor is generally not more 
than 10.degree. F. (5.degree. C.), and preferably not more than 2.degree. 
F. (1.degree. C.), less than the temperature of the outlet of the combined 
feed exchanger. Accordingly, the inlet temperature of the first reaction 
zone is generally from about 750.degree. to about 960.degree. F. 
(399.degree. to 516.degree. C.), preferably from about 800.degree. to 
about 900.degree. F. (427.degree. to 482.degree. C.). 
As mentioned previously, naphthene reforming reactions that are endothermic 
occur in the first reaction zone, and thus the outlet temperature of the 
first reaction zone is less than the inlet temperature of the first 
reaction zone and is generally from about 600.degree. to about 850.degree. 
F. (316 to 454.degree. C.). The first reaction zone contains generally 
from about 5% to about 50%, and more usually from about 10% to about 30%, 
of the total catalyst volume in all of the reaction zones. Consequently, 
the liquid hourly space velocity (LHSV) in the first reaction zone, based 
on the catalyst volume in the first reaction zone, is generally from 0.2 
to 200 hr.sup.-1, preferably from about 2 to about 100 hr.sup.-1, and more 
preferably from about 5 to about 20 hr.sup.-1. The catalyst particles that 
are withdrawn form the first reaction zone and passed to the second 
reaction zone generally have a coke content of less than 2 wt-% based on 
the weight of catalyst. 
The first reaction zone effluent stream is heated in a heater, such as a 
gas-fired, an oil-fired, or a mixed gas-and-oil-fired heater, of a kind 
that is well known to persons of ordinary skill in the art of reforming. 
The heater may heat the first reaction zone effluent stream by radiant 
and/or convective heat transfer. Commercial fired heaters for reforming 
processes typically have individual radiant heat transfer sections for 
individual heaters and a common convective heat transfer section that is 
heated by the flue gases from the radiant sections. Thus, in one 
embodiment of this invention the first reaction zone effluent stream 
passes first through the common convective section and then through one of 
the radiant sections, so that the two sections jointly perform the 
function of one heater. This heater is often referred to as an 
interheater, because it is located between two reaction zones, namely in 
this case the first and second reaction zones. The first reaction zone 
effluent stream leaves the interheater at a temperature of generally from 
about 900 to about 1040.degree. F. (482 to 560.degree. C.). Accounting for 
heat losses, the interheater outlet temperature is generally not more than 
10.degree. F. (5.degree. C.), and preferably not more than 2.degree. F. 
(10.degree. C.), more than the inlet temperature of the second reaction 
zone. Accordingly, the inlet temperature of the second reaction zone is 
generally from about 900.degree. to about 1040.degree. F. (482.degree. to 
560.degree. C.), preferably from about 980.degree. to about 1000.degree. 
F. (527.degree. to 538.degree. C.), and most preferably from about 
987.degree. to about 993.degree. F. (531.degree. to 534.degree. C.). The 
inlet temperature of the second reaction zone is usually at least about 
60.degree. F. (33.degree. C.) greater than the inlet temperature of the 
first reaction zone, and may be at least about 100.degree. F. (56.degree. 
C.) or even at least about 150.degree. F. (83.degree. C.) higher than the 
first reaction zone inlet temperature. The inlet temperature of the second 
reaction zone is generally from about 60.degree. to about 150.degree. F. 
(33.degree. to 83.degree. C.), and preferably from about 100.degree. to 
about 120.degree. F. (56.degree. to 67.degree. C.), greater than the inlet 
temperature of the first reaction zone. 
This invention is particularly applicable to those reforming processes 
where, if all of the reaction zones operated at the same inlet 
temperature, then the desired reformate octane could be achieved with 
inlet temperatures that are well below the maximum operating, or design, 
temperature of the reaction zone. By "well below", it is meant that the 
inlet temperature is at least 20.degree. F. (11.degree. C.), and 
preferably at least 30.degree. F. (17.degree. C.) less than the maximum 
operating, or design, temperature of the reaction zone. The desired 
reformate octane of the C.sub.5 + fraction of the reformate is generally 
from 85 to 107 clear research octane number (C.sub.5 + RONC), and 
preferably from 98 to 102 C.sub.5 + RONC. The explanation as to why this 
invention is particularly applicable to such a process relates to the fact 
that the inlet temperature of the first reaction zone is generally from 
about 60 to about 150.degree. F. (33 to 83.degree. C.) less than the inlet 
temperature of the second reaction zone. In the more general case where 
there is at least one additional interheater-reactor pair downstream of 
the second reactor, the inlet temperature to the first reaction zone is 
generally from about 60 to about 150.degree. F. (33 to 83.degree. C.) less 
than the inlet temperature of any other reaction zone downstream of its 
paired interheater. 
As persons skilled in the art of reforming know, the reformate product 
quality (C.sub.5 + RONC) in reforming units is generally correlated with a 
weight average of the inlet temperatures of the reaction zones, or of the 
average temperatures of the reaction zones. As used herein, the term 
"weight average temperature of the inlet temperatures of the reaction 
zones," which is commonly referred to as the "weight average inlet 
temperature" or "WAIT," is the sum of the products of the inlet 
temperature for each reactor and the weight of catalyst in that reactor, 
where the weight of catalyst in that reactor is expressed as a weight 
fraction of the total weight of catalyst in all of the reactors. 
Similarly, the term "weight average temperature of the average 
temperatures of the reaction zones," which is commonly referred to as the 
"weight average bed temperature" or "WABT," is the sum of the products of 
the bed temperature for each reactor and the weight of catalyst in that 
reactor, where the bed temperature is the arithmetic average of the inlet 
and outlet temperatures for each reaction zone and where, in the same 
manner as for WAIT, the weight of catalyst in each reactor is expressed as 
a weight fraction of the total weight of catalyst in all of the reactors. 
Thus, this invention is particularly applicable to those reforming 
processes where, if all of the reaction zones operated at the same inlet 
temperature, then the desired reformate octane could be achieved with 
inlet temperatures that are at least 20.degree. F. (11.degree. C.) less 
than the maximum operating, or design, temperature of the reaction zone 
because, in order to maintain the same WAIT or WABT when the inlet 
temperature of the first reaction zone is generally from about 60 to about 
150.degree. F. (33 to 83.degree. C.) less than the inlet temperature of 
the second reaction zone (and that of downstream reaction zones, if any), 
the inlet temperatures of the second reaction zone (and that of downstream 
reaction zones, if any) must be increased slightly. If the inlet 
temperatures of all of the reaction zones are at the same temperature and 
are already within 20.degree. F. (11.degree. C.) of their maximum 
operating, or design, temperatures, then it would not be possible as a 
practical matter to raise the inlets of those other reactors in order to 
maintain the same WAIT. Accordingly, the reformate C.sub.5 + RONC would 
fall. A temperature difference of at least 20.degree. F. (11.degree. C.) 
between the inlet temperature and the maximum operating, or design, 
temperature is desirable in order to provide operating flexibility in the 
event of catalyst deactivation. 
The second reaction zone contains generally from about 10% to about 60%, 
and more usually from about 15% to about 40%, of the total catalyst volume 
in all of, the reaction zones. Consequently, the liquid hourly space 
velocity (LHSV) in the second reaction zone, based on the catalyst volume 
in the second reaction zone, is generally from 0.13 to 134 hr.sup.-1, 
preferably from about 1.3 to about 67 hr.sup.-1, and more preferably from 
about 3.3 to about 13.4 hr.sup.-1. 
The second reaction zone effluent can pass to a train of pairs of heaters 
and reactors that contain moving catalyst beds. That is, the second 
reaction effluent can pass a second interheater (the first interheater 
being the previously described interheater between the first and the 
second reaction zones), and after heating, can pass to a third reaction 
zone. However, one or more additional heaters and/or reactors after the 
second reaction zone are not essential elements of this invention; that 
is, the second reaction zone may be the last reaction zone in the train. 
The third reaction zone contains generally from about 25% to about 75%, 
and more usually from about 30% to about 50%, of the total catalyst volume 
in all of the reaction zones. Likewise, the third reaction zone effluent 
can pass to a third interheater and from there to a fourth reactor. The 
fourth reaction zone contains generally from about 30% to about 80%, and 
more usually from about 40% to about 50%, of the total catalyst volume in 
all of the reaction zones. The inlet temperatures of the third, fourth, 
and subsequent reaction zones are generally within from about 18.degree. 
F. (10.degree. C.) of the inlet temperature of the second reaction zone. 
Because the reforming reactions that occur in the second and subsequent 
(i.e., third and fourth) reaction zones are generally less endothermic 
than those that occur in the first reaction zone, the temperature drop 
that occurs in the later reaction zones is generally less than that that 
occurs in the first reaction zone. Thus, the outlet temperature of the 
last reaction zone may be 20.degree. F. (11.degree. C.) or less below the 
inlet temperature of the last reaction zone, and indeed may conceivably be 
higher than the inlet temperature of the last reaction zone. This 
invention is most applicable when the temperature drop across the last 
reaction zone is relatively small or nonexistent, because as the outlet 
temperature of the last reaction zone increases the heat carried by the 
effluent of the last reaction zone effluent stream can be more readily 
transferred to the combined feed stream via the combined feed heat 
exchanger. In addition, for a given hot end approach in a countercurrent 
combined feed exchanger, as the temperature of the last reaction zone 
effluent stream increases, the temperature of the heated combined feed 
stream to the first reaction zone increases. 
As previously mentioned, the last reaction zone effluent stream is cooled 
in the combined feed heat exchanger by transferring heat to the combined 
feed stream. After leaving the combined feed heat exchanger, the cooled 
last reactor effluent passes to a product recovery section. Suitable 
product recovery sections are known to persons of ordinary skill in the 
art of reforming and do not, therefore, require detailed description 
herein. Briefly, such product recovery facilities generally include 
gas-liquid separators for separating hydrogen and C.sub.1 -C.sub.3 
hydrocarbon gases from the last reactor effluent stream, and fractionation 
columns for separating at least a portion of the C.sub.4 -C.sub.5 light 
hydrocarbons from the remainder of the reformate. In addition, the 
reformate may be separated by distillation into a light reformate fraction 
and a heavy reformate fraction 
The drawing illustrates an embodiment of the present invention. The drawing 
is presented solely for purposes of illustration and is not intended to 
limit the scope of the invention as set forth in the claims. The drawing 
shows only the equipment and lines necessary for an understanding of the 
invention and does not show equipment such as pumps, compressors, heat 
exchangers, and valves which are not necessary for an understanding of the 
invention and which are well known to persons of ordinary skill in the art 
of hydrocarbon processing. 
The drawing shows a common reaction vessel 100 that contains four stacked 
reaction zones: an upper first reaction zone 10, an intermediate second 
reaction zone 20, an intermediate third reaction zone 30, and a bottom 
fourth reaction zone 40. These four reaction zones are sized as to length 
and annular cross-sectional area of the catalyst bed such that the 
distribution of the total catalyst volume is 10% in reaction zone 10, 15% 
in reaction zone 20, 25% in reaction zone 30, and 50% in reaction zone 40. 
Although the reaction zones 10, 20, 30, and 40 are contained in a common 
reaction vessel 100, the reaction zones may be contained in two or more 
reaction vessels. The drawing does not limit the number of reaction zones 
that are contained in a single reaction vessel. 
In normal operation, fresh or regenerated catalyst particles are introduced 
through a line 24 and an inlet nozzle 26 into first reaction zone 10. The 
catalyst particles flow by gravity from first reaction zone 10 to second 
reaction zone 20, from second reaction zone 20 to third reaction zone 30, 
and from third reaction zone 30 to fourth reaction zone 40. The catalyst 
particles are ultimately withdrawn from common reaction vessel 100 through 
an outlet nozzle 52 and a line 54. Catalyst particles withdrawn through 
the line 54 may be transported to a conventional continuous regeneration 
zone, which is not shown in the drawing. The flow rate of catalyst through 
the common reactor vessel 100 can be controlled by regulating the rate of 
withdrawal of catalyst particles through line 54 in order to achieve a 
desired degree of catalytic performance (i.e., activity of catalyst, yield 
of desired products, and selectivity of desired products over undesired 
by-products) in the reaction zones 10, 20, 30, and 40. 
Turning next to the flow of hydrocarbons, a feedstock comprising a 
straight-run naphtha gasoline fraction boiling in the 180-400.degree. F. 
(82-204.degree. C.) range is charged to the process through a line 12 and 
is admixed with a hydrogen-rich gas stream flowing through a line 16 to 
form a combined feed stream. The combined feed stream flows through a line 
14 to a plate type combined feed heat exchanger 50, which heats the 
combined feed stream by heat exchange with the effluent stream of fourth 
reaction zone 40 flowing through a line 56. The heated combined feed 
stream passes through a line 22 and to the first reaction zone 10. 
An effluent stream is recovered from the first reaction zone 10 through a 
line 28 and becomes the feed stream to the second reaction zone 20. 
Because reforming reactions are generally endothermic, the second reaction 
zone feed stream passes through a heater 32 which reheats the stream to 
the desired inlet temperature of the second reaction zone 20. After 
heating, the second reaction zone feed stream passes through a line 34 to 
enter second reaction zone 20. An effluent stream is recovered from the 
second reaction zone 20 through a line 36. 
The effluent stream from the second reaction zone 20 passes through line 
36, through a heater 38 which heats the stream to the desired inlet 
temperature of the third reaction zone 30, and then through a line 42 to 
enter third reaction zone 30. Typical inlet temperatures and reaction 
pressures for the third reaction zone 30 are the same as those for the 
second reaction zone 20. An effluent stream is recovered from the third 
reaction zone 30 through a line 44. 
The effluent stream from the third reaction zone 30 passes through the line 
44, through a heater 46 which heats the stream to the desired inlet 
temperature of the fourth reaction zone 40, and then through a line 48 to 
enter fourth reaction zone 40. Typical inlet temperatures and reactor 
pressures for the fourth reaction zone 40 are the same as those for the 
second and third reaction zones, 20 and 30, respectively. An effluent 
stream is recovered from the fourth reaction zone 40 through the line 56. 
The effluent stream from the fourth reaction zone 40 passes to the plate 
type combined feed heat exchanger 50, which cools the effluent stream by 
heat exchange with the combined feed stream flowing through the line 14. 
The fourth reaction zone effluent stream then passes through a line 52 to 
a cooler 54 which cools the effluent stream to the desired inlet 
temperature of the separator 60, and then passes through a line 56 to 
separator 60. In separator 60, the effluent stream is separated into a 
hydrogen-containing gas stream that is withdrawn through a line 18 and a 
liquid stream containing the product reformate that is withdrawn through a 
line 58. One portion of the hydrogen-containing gas stream flows through a 
line 16, combines with straight-run naphtha being charged to the process, 
and is recycled to the common reaction vessel 100, as described 
previously. Another portion of the hydrogen-rich gas stream is passed 
through a line 62 to conventional product separation facilities, which are 
not shown in the drawing, for recovery of a hydrogen-rich gas stream. By a 
hydrogen-rich gas stream, it is meant a gas stream having a hydrogen 
content of at least 50 mol-%. The product reformate stream is passed 
through the line 58 to conventional product separation facilities, which 
are also not shown in the drawing, for recovery is of high octane product, 
for example, a reformate having a C.sub.5 + RONC of about 100. 
EXAMPLE 
The following example is intended to further illustrate the subject 
process. This illustration of an embodiment of the invention is not meant 
to limit the claims of this invention to the particular details of this 
example. This example is based on engineering calculations and actual 
operating experience with similar processes. 
This example compares the reactor inlet temperatures of a 
4-reactor/4-heater process and a 4-reactor/3-heater process. The two 
processes are both moving bed processes with continuous regeneration which 
each reform the same feedstock at the same feed rate. The LHSV, hydrogen 
to hydrocarbon molar ratio, reactor pressure, catalyst, C.sub.5 + RONC, 
and catalyst circulation rate each are the same in both processes. The 
catalyst distribution in the four reactors of each process is 
14%/19%/22%/45%. For the 4-reactor/4-heater process, the inlet temperature 
of each of the four reactors is 979.degree. F. (526.degree. C.), so the 
WAIT is 979.degree. F. (526.degree. C.). For the 4-reactor/3-heater 
process, the inlet temperature of the first reactor is 858.degree. F. 
(459.degree. C.), and the inlet temperature of each of the second, third, 
and fourth reactors is 991.degree. F. (533.degree. C.). Thus, the inlet 
temperature of each of the second, third, and fourth reactors is 
133.degree. F. (74.degree. C.) greater than the inlet temperature of the 
first reactor. For the 4-reactor/3-heater process, the WAIT is 972.degree. 
F. (522.degree. C.). Even though the reactor costs of the two processes 
are not significantly different, the 4-reactor/3-heater process has a 
lower capital cost due to its fewer number of heaters than the 
4-reactor/4-heater process, while producing the same quantity and quality 
of reformate.