Isomerization process with improved chloride recovery

A process for the isomerization of hydrocarbons using a chloride promoted catalyst wherein an adsorption zone arrangement operates to maintain chloride compounds in the reaction zone and to prevent contamination of product streams with chloride compounds removes normal paraffins from the desorbent stream to extend the capacity and life of a clinoptilolite molecular sieve. The invention preferably uses isoparaffins recovered from the isomerization zone as a desorbent. A gaseous fraction of the isomerization zone effluent can be recovered to provide a desorbent containing a low concentration of normal paraffins.

BACKGROUND OF THE INVENTION 
1. Field of the Invention 
This invention relates generally to the catalytic conversion of 
hydrocarbons and the recovery of chloride compounds. 
2. Description of the Prior Art 
Many isomerization processes employ a highly effective chlorided platinum 
alumina catalyst system in the reaction zone. The chlorided catalyst 
requires a continual addition of chloride to replace chloride lost from 
the surface of the catalyst into the product stream. Hydrogen chloride 
and/or volatile organic chlorides escape from the process with a 
stabilizer overhead stream and, apart from the loss of chloride, pose 
environmental concern. In addition to the loss of chlorides and 
environmental concerns, chloride loss hinders the operation of chloride 
promoted isomerization zones in other ways. For example, the recycle of 
hydrogen or hydrocarbons by a zeolitic adsorption process is not practical 
when a chloride type catalyst is used unless hydrogen chloride is removed 
from the recycle stream. Hydrogen chloride produced by the addition of 
chloride to the reaction zone or released from the catalyst composite 
results in significant amounts of hydrogen chloride leaving the effluent 
from the isomerization zone. Contact of this hydrogen chloride with the 
crystalline alumino-silicates in adsorption or conversion zones will 
decompose the matrix structure of many crystalline alumino-silicates 
thereby destroying any catalytic or adsorptive function. Therefore, absent 
chloride neutralization methods, chlorided catalyst systems generally have 
insufficient compatibility with many zeolitic catalysts or adsorbents to 
permit simultaneous use. 
U.S. Pat. No. 5,336,834 issued Zarchy et. al. discloses a chloride 
activated isomerization process that uses a thermal swing adsorption zone 
to recover hydrogen chloride from the isomerization zone and recycle the 
hydrogen chloride to the inlet of the isomerization zone. The process uses 
an acid tolerant molecular sieve to adsorb HCl from the stabilizer 
overhead and produce an HCl free effluent stream. The HCl is desorbed into 
the isomerization zone feed to completely contain the chloride compounds 
within the isomerization zone. Zarchy et. al teach the use of hydrocarbon 
feed as the desorbent to the isomerization zone. This feed is rich in 
normal hydrocarbons. It has been found that repeated adsorption and 
desorption of chloride compounds with isomerization zone effluent and feed 
streams diminish the chloride retention capacity of the acid resistant 
adsorbent. 
It is an object of this invention to improve the methods of recovering and 
recycling chloride compounds in an isomerization process that uses a 
chloride promoted isomerization catalyst. 
SUMMARY OF THE INVENTION 
It is now been discovered that the life of an acid resistant adsorbent can 
be extended indefinitely by using a desorbent that is substantially free 
of normal hydrocarbons. In particular, it has been discovered that the 
desorbent should be substantially free of normal hydrocarbons containing 
more than three carbon atoms. The effect of the normal hydrocarbons and, 
in particular, the heavier hydrocarbons is to cause a progressive and 
permanent deactivation of the adsorbent. Therefore, by the elimination of 
normal hydrocarbons from the desorbent stream, this invention can provide 
an indefinitely longer life to an acid resistant adsorbent. 
The benefits of this invention are further realized by the use of a 
specialized flow scheme for integrating the adsorption step with the 
isomerization process while essentially eliminating the passage of normal 
hydrocarbons as desorbent to an adsorbent containing adsorbed chloride 
compounds. For example, in a butane isomerization process, the flow scheme 
uses a deisobutanizer overhead product as a desorbent. The deisobutanizer 
overhead is substantially free of normal hydrocarbons. The selection of 
the deisobutanizer overhead stream permits desorption of the 
chloride-containing adsorbers with a substantially normal hydrocarbon-free 
stream. The flow scheme can also be arranged to prevent contacting of the 
adsorbent during the adsorption stage with heavy paraffins. 
Accordingly in one embodiment this invention is a process for the 
isomerization of a hydrocarbon feed containing normal hydrocarbons with a 
chloride promoted catalyst. The process combines a feed stream comprising 
normal paraffins with a first desorption stream comprising a chloride 
compound and containing less than 10 wt % normal paraffins to produce a 
combined feedstream. The combined feedstream contacts a chloride promoted 
isomerization catalyst at isomerization conditions to convert normal 
paraffins to isoparaffins and produce an effluent stream comprising normal 
paraffins, isoparaffins and a chloride compound. At least a portion of the 
effluent stream passes to an adsorption zone that contacts the portion of 
the effluent with an adsorbent to adsorb chloride compounds from the 
effluent stream. A second desorbent stream containing less than 10 wt % 
normal paraffins enters the adsorption zone to desorb chloride compounds 
from the adsorbent and produce the first desorption stream. Isoparaffins 
from the effluent stream are recovered from the process. 
In another embodiment this invention is a process for the isomerization of 
a hydrocarbon feed containing normal hydrocarbons with a chloride promoted 
catalyst that combines a feed stream comprising normal paraffins with a 
desorption stream comprising isoparaffins and a desorbed chloride compound 
to produce a combined feedstream. The combined feedstream contacts a 
chloride promoted isomerization catalyst at isomerization conditions to 
convert normal paraffins to isoparaffins and produce an effluent stream 
comprising normal paraffins, isoparaffins and a chloride compound. The 
effluent stream is passed to a first separator for recovery of an overhead 
stream containing the chloride compound and a bottoms stream comprising 
normal paraffins and isoparaffins. The overhead stream passes to an 
adsorption zone that contacts the overhead stream with a clinoptilolite 
molecular sieve to adsorb said chloride compound from the overhead stream. 
The bottoms stream passes to a second separator to separate normal 
paraffins from isoparaffins. An isoparaffin stream rich in isoparaffins is 
recovered from the separator. A portion of the isoparaffin stream passes 
to the adsorption zone to desorb chloride compounds from the desorbent and 
produce the desorption stream comprising isoparaffins and chloride 
compounds. A portion of the isoparaffin stream is recovered from the 
process as product. 
Additional embodiments and details of this invention are disclosed in the 
following Detailed Description of the Invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS 
The preferred embodiment of this invention is in the isomerization of 
C.sub.4 -C.sub.6 hydrocarbons. The products of isomerization processes 
contribute to a gasoline blending pool. Such gasoline blending pools 
normally include C.sub.4 and heavier hydrocarbons having boiling points of 
less than 205.degree. C. (395.degree. F.) at atmospheric pressure. This 
range of hydrocarbons includes C.sub.4 -C.sub.7 paraffins and especially 
the C.sub.5 and C.sub.6 normal paraffins which have relatively low octane 
numbers. The C.sub.4 -C.sub.6 hydrocarbons have the greatest 
susceptibility to octane improvement by lead addition and were formerly 
upgraded in this manner. Octane improvement is now often obtained by using 
isomerization to rearrange the structure of the straight-chain paraffinic 
hydrocarbons into branch-chained paraffins. 
Preferred feedstocks for isomerization are rich in normal paraffins having 
from 4 to 6 carbon atoms or a mixture of such normal paraffins. The term 
"rich" is defined to mean a stream having more than 50% of the mentioned 
component. Other useful feedstocks include light natural gasoline, light 
straight run naphtha, gas oil condensate, light raffinates, light 
reformate, light hydrocarbons and straight run distillates having 
distillation end points of about 77.degree. C. (170.degree. F.) and 
containing substantial quantities of C.sub.4 -C.sub.6 paraffins. The feed 
stream may also contain low concentrations of unsaturated hydrocarbons and 
hydrocarbons having more than 7 carbon atoms. The concentration of these 
materials should be limited to 10 wt. % for unsaturated compounds and 20 
wt. % for heavier hydrocarbons in order to restrict hydrogen consumption 
and cracking reactions. 
The isomerization of paraffins is generally considered a reversible first 
order reaction. The reaction is limited by thermodynamic equilibrium. The 
most common types of catalyst systems that are used in effecting the 
reaction are hydrochloric acid promoted aluminum chloride systems and 
supported aluminum chloride catalysts. The isomerization reaction zone 
typically contains a fixed bed of a chloride promoted isomerization 
catalyst. In the expectation that the feedstock will contain some olefins 
and, therefore, will undergo at least some cracking, the catalyst is 
preferably combined with an additional catalyst component that will 
provide a hydrogenation-dehydrogenation function. Preferably, this 
component is a noble metal of Group VIII of the periodic classification of 
the elements which are defined to include ruthenium, rhodium, platinum, 
osmium, iridium and palladium, with these specific metals being also known 
as the platinum group metals. The catalyst composition can be used alone 
or can be combined with a porous inorganic oxide diluent as a binder 
material. Other suitable binders include alumino-silicate clays such as 
kaolin, attapulgite, sepiolite, polygorskite, bentonite and 
montmorillonite, when rendered in a pliant plastic-like condition by 
intimate admixture with water, particularly when the clays have not been 
acid washed to remove substantial quantities of alumina. 
Of these chlorided catalyst systems a particularly preferred type of 
catalyst consists of a high chloride catalyst on an alumina base 
containing platinum. The alumina may be selected from various forms 
including an anhydrous gamma-alumina with a high degree of purity. The 
catalyst may also contain other platinum group metals. These metals 
demonstrate differences in activity and selectivity such that platinum is 
the preferred metal for use in such catalysts. The catalyst will contain 
from about 0.1-0.25 wt. % platinum. Other platinum group metals may be 
present in a concentration of from 0.1-0.25 wt. %. The platinum component 
may exist within the final catalytic composite as an oxide or halide or as 
an elemental metal. The presence of the platinum component in its reduced 
state has been found most suitable for this purpose. The catalyst also 
contains the chloride component. The chloride component termed in the art 
"a combined chloride" is present in an amount from about 2 to about 10 wt. 
% based upon the dry support material. The use of chloride in amounts 
greater than 5 wt. % have been found to be the most beneficial in these 
catalysts. 
There are a variety of ways for preparing a chlorided catalytic composite 
and incorporating a platinum metal and chloride therein. The method that 
has shown the best results impregnates the carrier material through 
contact with an aqueous solution of a water-soluble decomposable compound 
of the platinum group metal. For best results, the impregnation is carried 
out by dipping the carrier material in a solution of chloroplatinic acid. 
Additional solutions that may be used include ammonium chloroplatinate, 
bromoplatinic acid or platinum dichloride. Use of the platinum chloride 
compound serves the dual function of incorporating the platinum component 
and at least a minor quantity of the chloride into the catalyst. 
Additional amounts of the chloride must be incorporated into the catalyst 
by the addition or formation of aluminum chloride to or on the 
platinum-alumina catalyst base. An alternate method of increasing the 
chloride concentration in the final catalyst composite is to use an 
aluminum hydrosol to form the alumina carrier material such that the 
carrier material also contains at least a portion of the chloride. 
Chloride may also be added to the carrier material by contacting the 
calcined carrier material with an aqueous solution of the chloride such as 
hydrogen chloride. 
When a chlorided catalyst is used, operation of the isomerization zones 
often uses a small amount of a chloride promoter. The chloride promoter, 
typically an organic chloride, serves to maintain a high level of active 
chloride on the catalyst as low levels are continuously stripped off the 
catalyst by the hydrocarbon feed. The concentration of promoter in the 
reaction zone is maintained at from 30-300 ppm. The preferred promoter 
compound is carbon tetrachloride. Other suitable promoter compounds 
include oxygen-free decomposable organic chlorides such as 
propyldichloride, butylchloride, and chloroform to name only a few of such 
compounds. The need to keep the reactants dry is reinforced by the 
presence of the organic chloride compound which may convert, in part, to 
hydrogen chloride. As long as the process streams are kept dry, there will 
be no adverse effect from the presence of small amounts of hydrogen 
chloride. 
These chloride promoted catalysts are very reactive and can generate 
undesirable side reactions such as disproportionation and cracking. These 
side reactions not only decrease the product yield but can form olefinic 
fragments that polymerize or deposit on the catalyst and shorten its life. 
One commonly practiced method of controlling these undesired reactions has 
been to carry out the reaction in the presence of hydrogen. 
In accordance with this invention, an adsorption zone recovers the chloride 
that the feed strips from the catalyst. At least a portion of the effluent 
from the reaction zone contacts an adsorbent having a capacity for 
chloride compounds. Unless otherwise noted the term "portion" in this 
specification when describing a process stream refers to either an aliquot 
portion of the stream or a dissimilar fraction of the stream having a 
different composition than the total stream from which it was derived. 
The principle of the instant invention can be implemented with many 
different adsorbents such as zeolites, clays, inorganic polymers, alumina, 
silica gel, zirconia, carbon, organic polymers such as resin adsorbent, 
etc. Suitable adsorbents will have a selectivity and capacity for removal 
of chloride compounds in relatively low concentrations. Suitable 
adsorbents must also have an acid resistance to provide prolonged service 
in the chloride environment. Typically the chloride compounds will have a 
concentration of less than 3000 ppm in the effluent or effluent fractions 
from which the chloride compounds are adsorbed. The relatively small 
concentration of the chloride compounds in the isomerization effluent 
stream allow the process to use adsorbents have very low cyclic loadings. 
The process of this invention is viable with adsorbents that have cyclic 
loadings as low as 0.5 wt %. Preferably, the adsorbent material will 
recover 90 wt. % and, more preferably, more than 99 wt. % of the chloride 
compounds from the effluent or effluent fraction that it contacts. 
Adsorbents that are known to benefit the most when used in this invention 
are ion-exchanged clinoptilolites particularly, NaK clinoptilolite, acid 
washed Ba clinoptilolite, acid washed Na clinoptilolite, MgK 
clinoptilolite, acid leached clinoptilolite, NH.sub.4 clinoptilolite, etc. 
These adsorbents are particularly suitable for recovering and recycling 
HCl. Clinoptilolite as an adsorbent with adjustable pore size and acid 
resistance is described in U.S. Pat. No. 4,935,580 issued to Chao et al., 
U.S. Patent No. 4,964,889 issued to Chao, and U.S. Pat. No. 5,164,076 
issued to Zarchy, Chao and Correia, the contents of which are hereby 
incorporated by reference. For NaK and MgK clinoptilolite, the preferred 
concentration of potassium ions is in the range of 15 to 75%, and more 
preferably in a range of 30 to 70%, and the concentration of Na and Mg are 
preferably in a range of 25 to 85% and more preferably in a range of 30 to 
70%. The sum of Na and K or Mg and K are in the range of 50 to 100% of the 
total ion exchange capacity of the adsorbent. For Ba clinoptilolite, the 
barium concentration should be in the range of 20 to 100% of the ion 
exchange capacity. Other useful compositions include K, Na, Li, H, Mg, Ca, 
Sr, Zn, Mn, Co, CaK, SrK, ZnK, MnK, CoK and BaK cation exchanged or 
naturally occurring clinoptilolites and their acid washed version with the 
intended cations accounting for 50 to 100% of the total ion-exchange 
capacity of the clinoptilolite. 
The most important factor in obtaining suitable clinoptilolite materials 
for the adsorption of HCl compounds is the adjustment of the adsorbent 
pore size. In most cases, the changes in the pore size of zeolites 
following ion-exchange are consistent with a physical blocking of the pore 
opening by the cation introduced. In general, in any given zeolite, the 
larger the radius of the ion introduced, the smaller the effective pore 
diameter of the treated zeolite (for example, the pore diameter of 
potassium A zeolite is smaller than that of sodium A zeolite), as measured 
by the size of the molecules which can be adsorbed into the zeolite. 
Such is not the case, however, with clinoptilolites which demonstrate an 
unpredictable relationship that is not a simple function of the ionic 
radius of the cations introduced, i.e., pore blocking. For example 
potassium cations, which are larger than sodium cations, provide a 
clinoptilolite having a larger effective pore diameter than sodium 
ion-exchanged clinoptilolite. Sodium has an ionic radius of 0.98 .ANG. 
versus 1.33 .ANG. for potassium. See F. A. Cotton, G. Wilkinson, Advanced 
Inorganic Chemistry, Interscience Publishers (1980) or the Handbook of 
Chemistry and Physics, 56 Edition, CRC Press (1975) at pg. F-209, said 
references hereby incorporated by reference. In fact, a sodium 
ion-exchanged clinoptilolite with a sodium content equivalent to about 90% 
of its ion-exchange capacity defined by its aluminum content essentially 
excludes both hydrogen sulfide and n-butane. On the other hand, a 
potassium ion-exchanged clinoptilolite with a potassium content equivalent 
to about 95% of its ion-exchange capacity adsorbs hydrogen sulfide rapidly 
but substantially excludes n-butane. Thus, the clinoptilolite containing 
the cation with the larger ionic radius, i.e., potassium, has a larger 
pore than the clinoptilolite containing the cation with the smaller ionic 
radius, i.e., sodium. 
The clinoptilolites used in the process of the present invention may be 
natural or synthetic clinoptilolites. Natural clinoptilolites are 
preferred because they are currently readily available in commercial 
quantities. However, natural clinoptilolites are variable in composition 
and chemical analysis shows that the cations in clinoptilolite samples 
from various mines and even within a single deposit can vary widely. 
Moreover, natural clinoptilolites frequently contain substantial mounts of 
impurities, especially soluble silicates, which may alter the adsorption 
properties during activation, or may cause undesirable side effects which 
may inhibit practicing this invention. As an example of the compositional 
variations in natural clinoptilolites, the following Table 1 sets forth 
the chemical analysis of several clinoptilolite ore samples. 
TABLE 1 
______________________________________ 
Ore No. 1 2 3 4 5 
Source No. 
1 2 3 2 1 
______________________________________ 
Wt. % dry basis 
SiO.sub.2 
76.37 76.02 75.24 76.67 76.15 
Al.sub.2 O.sub.3 
12.74 13.22 12.62 13.95 12.90 
MgO 0.55 0.77 2.12 0.76 0.33 
CaO 0.55 2.19 2.72 2.27 1.04 
Na.sub.2 O 
3.86 3.72 2.25 3.26 4.09 
K.sub.2 O 
4.21 2.11 2.17 1.93 4.08 
Other* 1.72 1.98 2.88 1.16 1.41 
100.00 100.00 100.00 100.00 
100.00 
Elemental 
Concentration 
mmol/gm 
Si 12.73 12.67 12.54 12.78 12.69 
A1 2.50 2.59 2.47 2.74 2.53 
Mg 0.14 0.19 0.53 0.19 0.08 
Ca 0.10 0.39 0.49 0.41 0.19 
Na 1.25 1.20 0.73 1.05 1.32 
K 0.89 0.45 0.46 0.41 0.87 
______________________________________ 
*Includes the following oxides: Fe.sub.2 O.sub.3, SrO, BaO 
It can be seen from Table 1 that the concentrations of the various cations 
of the ore samples can vary quite substantially, especially when 
considered in view of the total theoretical ion-exchange capacity based on 
aluminum content. Note, for instance, the calcium content which varies 
from about 8 equivalent percent in Ore No. 1 to about 40 equivalent 
percent in Ore No. 3, e.g., for Ore No. 1, using the cation 
concentrations, Ca.times.2/Al.times.100=%, 
0.10.times.2/2.5.times.100=8.0%. Similarly, the potassium content varies 
from 15.0 equivalent percent in Ore No. 4 to 35.6 equivalent percent in 
Ore No. 1. With respect to cations present in relatively small amounts 
such as barium or strontium, the variations are generally not significant. 
Often, due to the above-described compositional variations, it is desirable 
to treat the natural clinoptilolite with a thorough ion-exchange to create 
a uniform starting material. For this initial ion-exchange, it is 
important to use a cation of reasonably high ion-exchange selectivity so 
it can effectively displace a substantial portion of the variety of 
cations originally existing in the natural zeolite. However, it is also 
important to not use a cation of overly high selectivity, otherwise it 
would make further tailoring of the adsorption properties of the 
clinoptilolite by ion-exchange difficult. The cations suitable to provide 
compositional uniformity in accordance with the present invention include 
sodium, potassium, calcium, lithium, magnesium, strontium, zinc, copper, 
cobalt, and manganese. It is often economically advantageous, and 
preferred, to use sodium or potassium for this purpose. The ion-exchanged 
clinoptilolite can then be further ion-exchanged with other cations, e.g., 
barium cations, to establish the desired level. It is, of course, possible 
to ion-exchange the clinoptilolite directly with cations other than those 
set forth above, e.g., barium cations, without an initial ion-exchange. 
Clinoptilolite typically loses some of its adsorption capacity for HCl 
after adsorbing and desorbing HCl. However, the rate of loss slows 
drastically after the first few cycles. It has been found that NaK 
clinoptilolite, acid washed Ba clinoptilolite, acid washed Na 
clinoptilolite, MgK clinoptilolite, acid leached clinoptilolite (or H 
clinoptilolite) and acid washed NH.sub.4 clinoptilolite all retain a 
substantial amount of HCl capacities after repeated adsorption and 
desorption in the absence of normal hydrocarbon desorbents. Another reason 
for the loss of HCl capacity after repeated cycles is the reaction of 
clinoptilolite cations with HCl to form chloride salt. In some cases, HCl 
washing to remove such chloride salt in the adsorbent manufacturing stage 
is helpful in providing a product that minimizes the phenomena of chloride 
salt formation and has a higher steady state HCl capacity. 
The type and concentration of hydrocarbons present in the feedstream have 
been found to have an important impact on the performance of the 
adsorbent. Hydrocarbons for the isomerization process will generally be 
present in a carbon range of from about 1 to about 8 carbon atoms per 
molecule. Propane and lighter hydrocarbons are often by-products produced 
by cracking in the isomerization zone. The adsorbent should have a low 
capacity for these small hydrocarbon impurities. NaK clinoptilolite, acid 
washed Ba clinoptilolite, and MgK clinoptilolite, as previously described, 
have suitably low capacity for C.sub.3 and lighter hydrocarbons and are 
preferred adsorbents. 
In accordance with this invention the loading of different normal paraffins 
will reduce the capacity of adsorbents for the cyclic adsorption of 
chloride compounds. Light paraffins having three atoms or less as well as 
heavier paraffins having four carbon atoms or more cause a loss in 
adsorption capacity for chloride compounds when present in repeated 
adsorption and desorption cycles. The nature of the deactivation differs 
depending on the type of light hydrocarbons. 
C.sub.3 and lighter paraffins cause a decrease in desorption capacity that 
is roughly in proportion to their concentration. However, the loading 
capacity loss caused by the light hydrocarbons in the adsorption feed have 
been found to be only temporary. 
The presence of heavier paraffins in a desorbent feedstream tend to cause a 
more gradual, but non-recoverable loss in chloride adsorption capacity 
after repeated adsorption and desorption cycles. The rate at which the 
adsorbent loses capacity for chlorides has been found to be proportional 
to the concentration of heavy normal paraffins present in the desorption 
stream. Decreases tend to occur linearly over the cycle life. For example, 
a desorption feed consisting of a fifty-fifty mixture of hydrogen and 
pentane linearly decreased the chloride adsorption capacity of a 
clinoptilolite adsorbent from 1.7 wt % to 0.3 wt % in 25 cycles when 
pentane portion was 100% normal pentane. A fifty-fifty mixture of pentane 
and hydrogen under the same adsorption and desorption conditions linearly 
decreased the adsorption capacity of the clinoptilolite adsorbent from 1.7 
to 1.2 wt % in the same number of cycles when the pentane portion was 50% 
isopentane and 50% normal pentane. The loading losses from desorption of 
chloride compounds in the presence of normal butane and heavier normal 
paraffins has been found to be non-recoverable by hot hydrogen stripping 
and vacuum activation. Air calcination at temperatures of up to 
370.degree. F. were found to offer some improvement in the hydrogen 
chloride loading capacity of an adsorbent. 
While not wishing to be bound by any theory, it is believed that the 
presence of the heavy normal paraffins creates surface pore clogging that 
makes the interior of the molecular sieve crystalline structure less 
accessible to the chloride compounds. Clogging of surface pores reduces 
the effective cyclic loading. Experiments have verified the pore clogging 
phenomenon by showing an increase in the mass transfer zone length after 
repeated cyclic loading using heavy normal paraffin desorbents. 
Suitable desorbent streams for this invention will have a concentration of 
C.sub.4 and heavier normal paraffins of less 10 wt %. More preferably the 
desorbent stream will have a concentration of C.sub.4 and heavier normal 
paraffins of less than 1 wt % and more preferably of less than 0.1 wt %. 
In a typical arrangement for the adsorption zone a fixed bed retains the 
adsorbent in the adsorption zone for contact with the input and effluent 
streams from the reaction zone. The adsorption zone preferably contains 
two or more adsorbent beds to continuously adsorb and desorb material from 
the effluent and the input stream. Typical conditions for operation of the 
adsorbent zones will again depend upon the particular adsorbents used and 
the temperature and pressure conditions of the inlet and effluent stream 
from the reaction zone. Typical conditions will include temperatures from 
50.degree.-750.degree. F. and pressures of from 1 atmosphere to 50 
atmospheres with the feedstream contacting the adsorbent in vapor or 
liquid phase conditions. Preferably, the processing conditions will 
maintain the feedstream through the adsorption zone in a vapor phase. The 
adsorption process will generally use temperature swing adsorption and 
desorption cycles with a typical temperature swing of from 100.degree. F. 
to 330.degree. F. 
Operating conditions within the isomerization zone are selected to maximize 
the production of isoalkane product from the feed components and are 
influenced by the type of catalyst as well as the composition of the feed. 
Two reaction zones are typically provided due to a temperature rise that 
initially occurs from hydrogenation reactions. Conditions within the first 
isomerization zone typically include a temperature in the range of 
190.degree.-290.degree. C. (375.degree.-550.degree. F.), a pressure of 
from 1200-3100 kPag (175-450 psig) and a liquid hourly space velocity of 
from 4-20. Typically, the reaction conditions are selected to keep the 
hydrocarbon feed in a vapor or mixed phase. Temperatures within the second 
conversion zone will usually operate at somewhat lower temperatures and 
range from about 65.degree.-280.degree. C. (150.degree.-536.degree. F.). 
These lower temperatures are particularly useful in processing feeds 
composed of C.sub.5 and C.sub.6 paraffins where the lower temperatures 
favor equilibrium mixtures having the highest concentration of the most 
branched paraffins. When the feed mixture is primarily C.sub.5 and C.sub.6 
paraffins, temperatures in the range of from 65.degree.-160.degree. C. 
(150.degree.-320.degree. F.) are preferred. When it is desired to 
isomerize significant amounts of C.sub.4 hydrocarbons, higher reaction 
temperatures are required to maintain catalyst activity. Thus, when the 
feed mixture contains significant portions of C.sub.4 -C.sub.6 paraffins 
most suitable operating temperatures are in the range from 
140.degree.-235.degree. C. (280.degree.-455.degree. F.). The second 
conversion zone may be maintained over the same range of pressures given 
for the first conversion zone. The feed rate to the second conversion zone 
may also vary over a wide range but will usually include liquid hourly 
space velocities that are lower than the first conversion zone and range 
from 0.5-12 hr..sup.-1, with space velocities of between 1 and 8 
hr..sup.-1 being preferred. The hydrogen concentration in the second 
conversion zone may also be adjusted by the addition of hydrogen to the 
feed or to the second conversion zone. The particular operating conditions 
within the isomerization zone will also be influenced by the makeup of the 
feed stream and the catalyst composition employed therein. 
Side reactions within the isomerization zone, particularly the saturation 
of unsaturates, will raise the temperature of the effluent from the first 
conversion zone. For example, the effluent from the first conversion zone 
can increase by 20.degree. F. for each percentage point of benzene that is 
present in the entering feed. As a remit of the increased temperature, the 
effluent from the first conversion zone is cooled in order to return it to 
a more desired isomerization temperature before it enters the second 
conversion zone. Even where there is not a substantial heat addition in 
the first conversion zone, it is often desirable to operate the second 
conversion zone in a two-stage isomerization process, at a lower 
temperature which, in the case of C.sub.5 -C.sub.6 hydrocarbons will move 
the reaction equilibrium toward the production of isoparaffins. The 
cooling is particularly beneficial for the arrangement of this invention 
where exothermic reactions can raise the temperature of the reaction zone 
effluent above those that are most beneficial for adsorption. 
Whether operating with one or more reactors the effluent from the 
isomerization zone will in most cases enter a separation zone for the 
removal of light gases from the isoparaffin containing product stream. The 
light gases include hydrogen added to the feed stream entering the first 
conversion zone and any additional hydrogen that was added to the feed 
entering the second conversion zone. At minimum, the separation facilities 
divide the conversion zone effluent into a product stream comprising 
C.sub.4 and heavier hydrocarbons and a gas stream which is made up of 
lighter hydrocarbons and hydrogen. Suitable designs for rectification 
columns and separator vessels are well known to those skilled in the art. 
The separation section may also include facilities for recovery of normal 
alkanes. Normal alkanes recovered from the separation facilities may be 
recycled to either the first or second conversion zone to increase the 
conversion of normal alkanes to isoalkanes. C.sub.3 and lighter 
hydrocarbons and any excess hydrogen from the second conversion zone are 
removed or returned to the process as part of the hydrogen gas stream. 
The type of separation zone and the hydrogen concentration in the effluent 
will influence the placement of the adsorption zone for the recovery of 
chloride compounds. Traditional isomerization processes operated with a 
relatively high recycle of hydrogen. In order to conserve hydrogen and 
stabilize the effluent, the effluent from the isomerization zone will pass 
directly to a hydrogen separator when the hydrogen to hydrocarbon ratio 
exceeds about 0.05. Cooling ahead of the hydrogen separator will lower the 
remainder of the effluent stream to a temperature in a range of from 
80.degree.-140.degree. F. Thus, the temperature condition of the effluent 
leaving the hydrogen separator is suitable for adsorption of chloride 
compounds. Where there is a hydrogen recycle, the chloride adsorber bed 
undergoing the adsorption step will normally receive the effluent stream 
passing from the hydrogen separator to a stabilizer for the separation of 
isomerate product from non-condensibles. Isomerization zone processes that 
operate with very low hydrogen concentrations in the feed and the effluent 
eliminate the hydrogen separator and an accompanying recycle compressor to 
reduce utility and capital cost. In these cases, the effluent stream 
passes directly from the isomerization reaction zone to the stabilizer. In 
most cases, the temperature of the effluent passing from the isomerization 
reaction zone to the stabilizer exceeds a suitable adsorption temperature. 
Therefore, where the isomerization reaction zone operates without hydrogen 
recycle, or in what is generally referred to as a hydrogen once-through 
operation, the effluent from the isomerization zone is cooled to place the 
chloride adsorber between the reactor and the stabilizer or the chloride 
adsorber generally adsorbs chloride compounds from the overhead of the 
stabilizer. Since most of the chloride compounds are contained in the 
stabilizer overhead, the isomerate from the stabilizer is still relatively 
free of chloride compounds and a majority of the chlorides are recovered 
for return to the isomerization reaction zone. Treating the overhead of 
the stabilizer has the advantage of virtually eliminating the presence of 
C.sub.4 and higher hydrocarbons from the adsorbent bed even during the 
adsorption cycle. 
Again, it is generally known that high chlorided platinum-alumina catalysts 
of this type are highly sensitive to sulfur and oxygen-containing 
compounds. A sulfur concentration of 0.5 ppm in the feed or less is 
required, since the presence of sulfur in the feedstock serves to 
temporarily deactivate the catalyst by platinum poisoning. Activity of the 
catalyst may be restored by hot hydrogen stripping of sulfur from the 
catalyst composite or by lowering the sulfur concentration in the incoming 
feed to below 0.5 ppm so that the hydrocarbon will desorb the sulfur that 
has been adsorbed on the catalyst. Water and oxygenate compounds are 
generally kept to a concentration of 0.1 ppm or less. The more stringent 
limitation on water and oxygenate compounds that decompose to form water 
stems from the fact that water can act to permanently deactivate the 
catalyst by removing high activity chloride from the catalyst and 
replacing it with inactive aluminum hydroxide. 
Although sulfur compounds will not cause permanent deactivation, 
isomerization feeds will usually contain sulfur which will interfere with 
the isomerization operations. Sulfur contaminants are present with the 
original crude oil fraction and include mercaptans, sulfides, disulfides 
and thiophenes. For light straight run feeds, sulfur concentrations will 
usually range from 20-300 ppm. Rapid coking of the catalyst has been 
experienced in most cases following sulfur deactivation. If left 
unchecked, the coking will be severe enough to require a complete 
regeneration of the catalyst. Therefore, it is common practice to minimize 
the amount of sulfur that contacts catalyst in the isomerization zone to 
prevent deactivation and avoid a full regeneration of the catalyst. 
A wide variety of adsorbents can be used for removing the sulfur from the 
feed. Suitable adsorbents for hydrogen sulfide include those adsorbents 
having a pore diameter of at least 3.6 .ANG., the kinetic diameter of 
hydrogen sulfide. Such adsorbents include Zeolite 5A, Zeolite 13X, 
activated carbon and other materials that are well known in the art and 
conventionally used for hydrogen sulfide adsorption. Particularly 
preferred adsorbents for the removal of hydrogen sulfide from the 
isomerization zone feedstream include 4A Zeolite and clinoptilolite 
molecular sieves. Zeolite 4A is the sodium form of Zeolite A and has pore 
diameters of about 4 .ANG.. The method for its preparation and chemical 
and physical properties are described in detail in U.S. Pat. No. 2,882,243 
the contents of which are hereby incorporated by reference. 
In order to demonstrate the preferred arrangement of this invention, an 
isomerization reaction zone using a chlorided platinum alumina catalyst 
and adsorptive chloride recovery section is depicted in FIG. 1. FIG. 1 is 
a schematic representation of the process and shows only the portions of 
the major equipment necessary to carry out the process. Other related 
equipment such as separators, pumps, compressors, coolers, condensers, 
reflux drums etc. which are well known to those skilled in the art have 
not been shown and are not necessary for an understanding of Applicant's 
invention or the underlying concepts. 
In the isomerization process of FIG. 1, a hydrocarbon feedstream containing 
principally normal butane and lesser amounts of propane and pentanes 
enters the process through a line 10. A desorption stream 12 containing 
HCl and isobutane receives hydrogen from a hydrogen containing gas stream 
14 to provide a stream 16 containing chloride and hydrogen that mixes with 
the feed. Hydrogen is added via line 14 in amounts that will produce a 
hydrogen to hydrocarbon ratio of less than 0.05. Preferably, the 
hydrogen-containing gas stream will have a hydrogen concentration greater 
than 75 wt. % hydrogen. A small amount of make-up chlorides are added to 
the line 16 via chloride input stream 15. A recycle stream 18 combines 
normal butane with the feed and stream 16 to provide a combined feed 20 
that enters a butane isomerization zone 22. Reaction zone 22 schematically 
represents an isomerization zone that can consist of multiple reactors, 
feed exchangers and heaters The isomerization zone contains chlorided 
platinum alumina catalyst that contacts the feedstream and produces an 
isomerization zone effluent stream 24. 
The effluent stream carried by line 24 enters a stabilizer zone 26 that 
separates lighter hydrocarbons and gases from heavier hydrocarbons in the 
entering effluent stream. A line 30 recovers C.sub.3 and lighter 
hydrocarbons overhead after it passes through ordinary parts of a reflux 
system such as a condenser and an overhead drum which are contained within 
the stabilizer zone 26. Stabilizer 26 separates a bottoms stream 28 
containing C.sub.4 and heavier hydrocarbons and minor amounts of C.sub.3 
hydrocarbons from the isomerization zone effluent. Line 30 recovers the 
remainder of the isomerization zone effluent as a net overhead stream from 
the stabilizer that contains less than 2 wt % C.sub.4 and higher normal 
paraffins. 
The net overhead stream 30 passes through a series of adsorption vessels 
within an adsorption zone 32 that contains as the adsorbent a 
clinoptilolite molecular sieve selected for the removal of chloride 
compounds from the overhead stream and resistance to the corrosive effects 
of the chloride compounds. Passage through the series of adsorbers within 
adsorption zone 32 produces an essentially chloride-free net overhead 
stream that contains C.sub.3 and lighter material. Adsorption vessels 
within adsorption zone 30 operate in cyclic fashion. A line 34 recovers a 
substantially chloride-free stream from adsorption zone 32. Stream 34 may 
be further processed for the recovery of hydrocarbons and gaseous 
components or, is more typically used as a fuel gas stream. 
Hydrocarbon liquid comprising the main product stream from the 
isomerization zone exits the stabilizer zone through line 28. Line 28 will 
typically receive liquid from a reboiler arrangement that is contained 
within the stabilizer zone 26. The stabilizer bottoms stream passes to a 
deisobutanizer column 36 that produces an isobutane rich stream 40, a 
bottoms stream of C.sub.5 and heavier hydrocarbons and normal butane 
recycle stream 18. 
A line 42 withdraws a small portion of an isobutane rich gas stream from 
the net deisobutanizer overhead 40 for desorbing chloride containing 
adsorbent in the adsorption zone 32. The isobutane stream contacts 
chloride containing adsorbent in an adsorber vessel that was formerly in 
an adsorption mode. The amount of isobutane needed from the net isobutane 
overhead stream for desorbing chloride compounds from the adsorber beds of 
adsorption zone 32 is less than 10% of the net overhead from the 
deisobutanizer and, more typically less than 2% of the net overhead. The 
chloride containing desorption fluid exits adsorption zone 32 via line 12 
and supplies the majority of the chloride needed for the promotion of the 
isomerization catalyst in reactor zone 22. Any small chloride losses are 
replaced by the additional chloride compounds from line 15. 
EXAMPLES 
The following examples were run to determine what effect the presence of 
normal paraffins had on the continued adsorption activity of 
clinoptilolite adsorbents. A series of tests were run in a pilot plant 
containing interchangeable adsorbers for the testing of one adsorbent 
loaded adsorber at a time. The adsorber supported a bed approximately 5/8 
inch in diameter by approximately 93/4 inches long. Chromatographic 
columns were arranged to measure HCl content in the adsorption and 
desorption effluent from the adsorbers. The tests used experimental 
conditions of 100.degree. F. and 214 psia for adsorption and 330.degree. 
F. and 214 psia for desorption. In each experiment NaK clinoptilolite was 
contained in each bed. 
Example I 
An adsorber containing a 28.5 gram sample of NaK clinoptilolite adsorbent 
which had been subjected to two passes of calcination at a temperature of 
600.degree. F. was cyclically contacted with a feed and a desorbent. The 
feed consisted of hydrogen containing 1798 PPMV hydrogen chloride. The 
feed flow rate through the adsorption step was approximately 800 cm.sup.3 
/min. Time until breakthrough was recorded for each adsorption cycle and 
the total weight percent loading at breakthrough for each cycle was 
recorded. Desorption was carried out for one hour in each cycle using a 
fifty-fifty mixture of normal pentane and isopentane at a flow rate of 1 
cm.sup.3 /min pentanes. Methane and ethane were excluded from these 
samples to avoid co-adsorption in order more readily to distinguish trends 
in the adsorption capacity of the adsorbent. As shown in FIG. 2, the 
loading steadily declined from an initial loading of 1.6 wt % HCl to 0.46 
wt % in 28 cycles. These test results established that the adsorbent was 
not commercially practical since the presence of methane and ethane--if 
added--would cause a further decrease in the adsorption capacity of the 
adsorbent. 
Example II 
Another experiment was conducted to confirm that normal pentane was causing 
the deactivation. This experiment used the identical feed conditions and 
adsorption procedure as Example I. This experiment differed from the 
previous pentane desorption experiment and used a fifty-fifty mixture on 
molar basis of pentane and hydrogen as a regenerant. The pentane again 
comprised a fifty-fifty mixture of isopentane and normal pentane resulting 
in a total normal pentane concentration of 25% on a volume basis. As FIG. 
2 demonstrates, the adsorbent immediately began deactivating upon addition 
of the pentane with a decrease in capacity from 1.77 wt % to about 1 wt % 
within 30 cycles. It is believed that the partial pressure of the hydrogen 
suppressed coke formation or the amount of pore clogging thereby 
decreasing the rate of the deactivation. However, there is no indication 
that the deactivation was leveling off. Therefore, although the addition 
of the hydrogen decreased the deactivation rate by about a factor of 2, it 
is not sufficient to arrest deactivation. 
Example III 
Another experiment was performed using a normal pentane free isopentane as 
a regenerant. Adsorption and desorption were again carried out in cycle 
steps of 4 hours and 1 hour, respectively. The feed during the adsorption 
step consisted of a hydrogen stream with 2200 PPMV HCl at a total flow 
rate of 772 cc/min. The adsorber apparatus retained approximately 34 grams 
of a NaK clinoptilolite prepared in substantially the same manner as that 
used in Example I. The desorbent for the first 27 cycles was a pure 
isopentane that passed through the adsorption zone at a flow rate of 1 
cc/min. As shown in FIG. 2, over the course of 27 cycles, the adsorbent 
provided a stable loading capacity of not less than 2.1 wt % HCl. 
Example IV 
The adsorption capacity loss and recovery due the presence of methane and 
ethane in the adsorption feed was investigated by charging 795 cc/min of a 
feedstream containing 250 cc/min of C.sub.1 -C.sub.2 and 1823 PPMV of HCl 
to an adsorber containing 29.38 grams of a sodium and potassium exchange 
clinoptilolite adsorbent. The adsorption and desorption cycles were 
essentially the same as those previously described except that the 
desorbent consisted of hydrogen at a flow rate of 250 scc/min. The initial 
flow rate to the adsorption zone consisted of 45 cc/min of HCl and 761 
cc/min of hydrogen. At about 18 cycles, about 250 cc/min of the hydrogen 
flow was replaced with a C.sub.1 -C.sub.2 mixture. After about 30 cycles, 
the C.sub.1 -C.sub.2 flow with the feed to the adsorption stream was again 
replaced with hydrogen. FIG. 3 dramatically illustrates the immediate 
effect of the methane and ethane to cause a sudden drop in HCl capacity. 
Just as dramatically, FIG. 3 illustrates how the removal of the C.sub.1 
-C.sub.2 from the feed caused an immediate return to the original capacity 
of the adsorbent. This testing illustrates that light paraffins have the 
immediate effect of dramatically reducing adsorption capacity for HCl of 
the adsorbents. This reduced capacity is immediately recovered upon 
removal of the methane or ethane. 
These examples demonstrate that the elimination of heavy normal paraffins 
from the desorbent stream of a chloride adsorption zone will prevent 
progressive deactivation of the adsorbent and the loss of the adsorbent's 
capacity for adsorption of chloride compounds. The data also demonstrates 
that light paraffins create temporary deactivation of adsorption capacity 
of the adsorbents for chloride compounds.