Note: Descriptions are shown in the official language in which they were submitted.
"ADSORPTIVE SEPARATION OF ISOPENTANE AND DI-METHYL
BRANCHED PARAFFINS FROM MONO-METHYL BRANCHED PARAFFINS"
~ . ~ . . . . .
Field of the Invention
This invention relates to processes for
separating by adsorption lower octane normal and
mono-methyl paraffins from hydrocarbon feeds
containing normal, mono-methyl and mGre highly
branched paraffin fractions, and isomerizing the
normal and mono-methyl paraffins to produce higher
octane isopentane and more highly branched paraffins.
Backqround of the Invention
Light straight run or virgin naphtha is a
hydrocarbon ref inery process stream comprising
pentane and hexane paraffins and is useful as a
blending stock for in gasoline pools. However, the
Research Octane Number of this hydrocarbon fraction
is relatively low, generally in the range of 68-72.
In the past, the octane of this fraction was
conveniently raised to about 85-90 RON by the
addition of alkyl lead compounds. More recently due
to lead phase-down, refiners have implemented other
means such as isomerization and reforming to improve
the octane of this fraction. Isomerization
processes typically produce a product having an
octane of about 78-84 RON depending upon the
temperature of the reaction. When the isomerization
processes are integrated with separation processes
such as adsorption or distillation, which separate
the effluent from the isomerization reactor into
higher and lower octane segments, the final product
from the integrated process typically has an octane
of about 90 RON. In acco~dance with the present
invention processes are provided for the adsorption
and isomerization of a hydrocarbon feed comprising
pentane and hexane normal and mono-methyl paraffins
which can produce a product having an octane o~
about 93-96 RON.
Molecular sieve adsorbents have been
utilized in a variety of processe in the
hydrocarbon processing industry. One area of
particular importance is in the field of octane
upgrading, where hydrocarbon feedstocks containing
pentane and hexane paraffin fractions are separated
into high and low octane se~ments. In general, the
normal paraffins and mono-methyl branched paraffins
comprise the low octane segment and the more highly
branched paraffins (e.g., di-methyl paraffins),
naphthenes, and aromatics comprise the high octane
segment. However, isopentane, which has a high
octane, also has a structure similar to that of the
other mono-methyl branched paraffins. Accordingly,
the adsorption processes currently existing in the
art cannot conveniently separate isopentane and
other high octane more highly branched paraffins
from the lower octane normal and mono-methyl
r paraffins, and as a result valuable high octane
components may be lost when upgrading the octane of
a hydrocarbon feed by adsorption. In addition to
the adsorption or separation function, most of the
processes used for octane upgrading also incorporate
an isomerization process that is used to further
isomerize the low octane normal and mono-methyl
2 ~
paraffins to higher octane isopentane and more highly
branched paraffins. Some or all of the effluent fro~ the
isomerization process can be recycled back to the adsorption
process for separationO Alternately, it can be combined
with effluent from the adsorption process to form a
combined total product.
The adsorption processes known in the art are
generally of two types. One type performs a separation of
normal from non-normal paraffins using an adsorbent ~ommonly
known in the industry as 5A or calcium zeolite A. ~his
process is useful because it can process a hydrocarbon feed
containing pentane and hexane paraffin fractions. While
this type of process provides substantial benefits, the
improvement in octane rating of the product is limited due
to the presence in the non-adsorbed fraction of low octane
mono-methyl paraffins which are not readily adsorbed by the
5A zeolite.
The other type of adsorption process incorporates
an adsorbent that has a slightly larger pore size which
allows both normal paraffins and mono-methyl paraffins to be
adsorbed but excludes the larger di-methyl branched
paraffins. U.S. Patent No. 4,717,784, e.g., at col. 3,
lines 59, et seq., describes an adsorption and isomerization
process that upgrades the octane of a C6 (hexane) paraffinic
feed by isomerizing the feed and subsequently separating the
unreacted normal paraffins and mono-methyl branched
paraffins from the di-methyl branched paraffins. This
process, however, fails to make any separation of the
relatively high octane mono-methyl paraffins having not more
than five carbon atoms, e.g., isopentane from the other
absorbed hexane and heavier mono-methyl paraffins, e.g. 2-
methyl pentene, and normal paraffin species. All absorbed
mono-methyl paraffins are desorbed along with the normal
paraffins and recycled to the isomerizer.
It can be appreciated that in light of the two
types of processes described above, processes are sought
~4~ 2 ~ 2 ~ 3
which can upgrade the octane of a hydrocarbon feedstoc~
containing pentane and hexane paraffinic fractions by
separating the low octane normal and mono-methyl paraffins
from isopentane and the higher octane, more highly branched
paraffinsO
Summary of the Invention
This invention broadly provides processes for the
separation of hydrocarbon feeds containing normal paraffins
and mono-methyl branched paraffins for purposes of octane
improvement. One aspect of this invention pertains to
processes for separating isopentane and di-methyl branched
paraffins from a hydrocarbon feed containing isopentane,
mono-methyl and di-methyl branched paraffins by passing said
hydrocarbon feed through an adsorber bed containing a
microporous molecular sieve adsorbent which has an
elliptical cross section with pore dimensions between 5 and
5.5 Angstroms along the minor axis and between 5.5 and 6
Angstroms along the major axis, e.g., silicalite at
adsorption conditions selected such that mono-methyl
branched paraffins including isopentane, are adsorbed and
isopentane is preferentially desorbed during continued
adsorption to provide a mass transfer zone having isopentane
concentrated at the leading edge thereof; removing an
essentially non-adsorbed fraction comprising di-methyl
branched paraffins from said adsorber bed as a first portion
of an adsorption effluent; eluting at least a portion of
the mass transfer zone which comprises isopentane from said
adsorber bed as a second portion of the adsorption effluent;
and thereafter desorbing the mono-methyl branched paraffins
from the adsorber bed under desorption conditions selected
to regenerate the bed and to produce a desorption effluent
comprising mono-methyl paraffins.
~5- 2~2~93
In a further aspect of this embodiment~ at
least a portion of the hydrocarbon feed is
pretreated to remove normal paraffins by passing the
hydrocarbon feed through an adsorber bed containing
adsorbent having pore dimensions sufficient to allow
adsorption of normal paraffins while essentially
excluding larger molecules such as mono-methyl
paraffins, e.g. calcium zeolite A.
In s~ill a further aspect of this
embodiment, at least a portion of the hydrocarbon
feed is obtained from an isomerization reactor and
at least a portion of the desorption effluent is
recycled to the isomerization reactor to form an
isomerization reactor effluent comprising normal,
mono-methyl branched, and more highly branched
paraffins.
Another aspect of this invention pertains
to processes for separating isopentane from a
hydrocarbon feed containing a pentane fraction which
comprises the steps of: passing said hydrocarbon
feed through an adsorber bed containing a
microporous molecular sieve having eliptical pores
with adsorbent cross sectional pore dimensions in
the range of approximately 4.5 to 5.5 Angstroms,
e.g. ZSM-23, such that isopentane and normal
pentane are adsorbed and isopentane is
preferentially desorbed during continued adsorption
-6- 2 0 2 !~
to provide a mass transfer zone having isopentane
concentrated at the leadi~g edge thereof; recovering
an adsorption effluent and eluting at least a
portion of the mass transfer zone which comprises
isopentane from said adsorber bed as adsorption
effluent; and desorbing adsorbed components as
desorption effluent.
Detailed DescriPtion of the Inventio_
The fresh feed treated by this process
contains normal, mono-methyl, and more highly
branched paraffins. It is composed principally of
the various isomeric forms of saturated hydrocarbons
having from five to about eight, preferably five to
six, carbon atoms. Often, the hydrocarbon feed
contains at least 40, most freguently 40 to 95 or
more, weight percent of such saturated
hydrocarbons. The expression "the various isomeric
forms" is intended to denote all the branched chain
and cyclic forms of the noted compounds, as well as
the straight chain forms. Also, the pr fix notations
"iso" and "i" are intended to be generic
designations of all branched chain and cyclic (i.e.,
non-normal) forms of the indicated compound unless
otherwise specif:ied.
~ he following composition is typical of a
feedstock containing 5 to 6 carbon atoms which is
suitable for processing according to the invention:
7 2~ ~13; ,? C,~
Components Mole %
C4 and lower o - 7
i-C5 10 - 40
n-C5 5 40
i-C6 10 - 40
n C6 5 - 30
C7 and higher 0 - 10
Suitable feedstocks containing 5 to 6
carbon atoms are typically obtained by refinery
distillation opera~ions, and may contain small
amounts of C7 and even higher hydrocarbons, but
these are typically present, if at all, only in
trace amounts. Olefinic hydrocarbons are
advantageously less than about 4 mole percent in the
feedstock. Aromatic and cycloparaffin molecules
have a relatively high octane number. Accordingly,
the preferred feedstocks are those high in aromatic
and cycloparaffinic hydrocarbons, e.g., at least 3,
and more typically from 5 to 25 mole percent of
these components combined.
In a 2referred aspect, the non-cyclic C5
and C6 hydrocarbons typically comprise at least
60, and more typically at least 75, mole percent of
the feedstock, with at least 25, and preferably at
least 35, mole percent of the feedstock being
hydrocarbons selected from the group of isopentane,
iso-hexane and combinations of these. Preferably,
the feedstock will comprise no more than 60, and,
-~- 2 ~ 2 ~
more preferably, no more than 50 mole percent of a
combination of n-pentane and n-hexane.
The processes of the present invention provide for
the separation of the low octane normal and mono-methyl
paraffins from the higher octane more highly branched
paraffins and isopentane by adsorption. The molecular sieve
used for adsorption in this process must be capable of
adsorbing normal paraffins as well as mono-methyl paraffins,
for instance, 2-methylpentane and 3-methylpentane, while
excluding larger di-methyl branched paraffins. Adsorbents
that are suitable for this application are those microporous
molecular sieves having pores of an elliptical cross section
with pore dimensions between about 5 and 5.5 Angstroms along
the minor axis and between about 5.5 and 6 Angstroms along
the major axis. A preferred adsorbent meeting this
specification is silicalite. Silicalite, as the term is
used herein includes both the silicapolymorph disclosed in
U.S. Patent No. 4,061,724 and also the F-silicalite
disclosed in U.S. Patent No. 4,073,865. Other suitable
adsorbents include ZSM-5, ZSM-11, ZSM-48, and other similar
crystalline aluminosilicates. ZSM-5 and ZSM-ll are
described in U.S. Patent No. 3,702,886 and Re. 29,948 and
U.S. Patent No. 3,709,979, the teaching of said patents
being incorporated herein by reference. In general these
adsorbents are high in silica content, the silica to alumina
~atio being at least 10:1 and the preferred adsorbents are
more commonly characterized as having silica to alumina
ratios higher than 100:1.
The adsorption section of the present invention
preferably contains at least two fixed beds of
~,?,~;3
_9_
adsorbent that cyclically undergo stages of
adsorption and desorption. Durinq the adsorption
step, or steps, feed containing normal pentane;
isopentane; 2,3 dimethylbutane (23DMB) ; 2,2
dimethylbutane (27DMB); 2-methylpentane (2MP);
3-methylpentane (3MP) and normal hexane is typically
introduced into the feed end of a bed containing the
adsorbent, e.g., silicalite. (For purposes of
discussion and ease of reference, the description is
specifically directed to typical C5 to C6
hydrocarbon feed. This is not intended to be in
limitation at the broad aspects of the invention.)
As the adsorption step progresses, 2,2
dimethylbutane and 2,3 dimethylbutane which are
essentially nonadsorbable, pass through the bed and
are discharged from the discharge end of the bed.
The other feed components are adsorbable to
different degrees and form adsorption fronts, or
mass transfer zones. The mass transfer zones travel
through the bed in an order inverse to their
relative adsorptivities. In other words, the less
strongly adsorbed components, e.g. isopentane, are
preferentially desorbed as more strongly adsorbed
components, such as n-pentane, are adsorbed during
continued adsorption. Hence, isopentane, being a
lesser strongly adsorbed component, is concentrated
at the leading edge of the mass transfer zone. An
exper;ment performed with a
silicalite adsorbent at a typical adsorption
temperature of 260C demonstrated the following
elution order.
. _ .
1 o 2 ~ 2 `q ~.t ?,~ 3
Component Breakthrough Time, Min.
. _ ~
22DMB 0.44
23DMB 0.88
iC5 3.10
nC5 4 79
3MP 6.14
2MP 7.05
nC6 12.05
- The foregoing experiment illustrates that
although isopentane is adsorbed by the silicalite,
it is less strongly adsorbed than normal pentane,
3-methylpentane, 2 methylpentane and normal hexane
and, therefore, elutes faster. Accordingly, the
adsorption step, as advantageously practiced
according to this invention, can be continued so as
to elute at least a portion of the mass transfer
zone, which comprises isopentane, from the adsorber
bed as an adsorption effluent.
In general, an~here from about 10 to
essentially loO weight percent of the isopentane
contained in the hydrocarbon feed processed
according to this invention can be eluted from the
adsorber bed as an adsorption effluent. Preferably,
however, at least about 50 weight percent, and, most
preferably, at least about 80 weight percent of the
the isopentane is eluted from the adsorber bed.
Upon completion of the adsorption step, the
adsorbent must undergo one or more desorption steps
to regenerate the adsorbent. Desorption may be
2 ~ ~3
accomplished by any conventional means including non-
adsorbable purge, displacement purge, temperature swing, and
pressure swing. A preferred method for regenerating the
adsorbent is by countercurrent desorption with a non-
adsorbable purge gas such as hydrogen. The non-adsorbable
purge gas is passed through the adsorber bed and is
continued until a major portion of the normal and mono-
methyl paraffin components have been desorbed and the bed
void space consists principally of non-adsorbable purge gas.
In a preferred embodiment, the adsorber feed is
pretreated to remove the normal paraffins prior to
adsorption of the mono-methyl paraffins and isopentane.
Adsorbent useful for such pretreatments can be described as
having pore dimensions sufficient to allow adsorption of
normal paraffins while essentially excluding larger
molecules such as mono-methyl and di-methyl branched
paraffins. Particularly suitable zeolites of this type are
the A-type zeolites, described in U.S. Patent No. 2,883,243,
which in several of its divalent exchanged forms, notably
the calcium cation form, have apparent pore diameters of
about 5 Angstroms, and have very large capacities for
adsorbing normal paraffins, while excluding mono-methyl
paraffins. Other suitable molecular sieves include zeolite
R, U.S. Patent No. 3,030,181; zeolite T, U.S. Patent No.
2,950,952, and the naturally occurring zeolitic molecular
sieves chabazite and erionite.
-12- ~2~ 3
In this particular aspect of the invention, the pre-
treatment adsorber bed containing, preferably, calcium
exchanged zeolite A is configured upstrea~,, preferably
immediately upstream, of the adsorber bed containing
a larger pore sized adsorbent, e.g. silicalite. The
two adsorbents can be contained in the same adsorber
vessel. It can be appreciated that when the two
adsorbents are utilized, the calcium zeolite A
adsorber bed can be efficiently desorbed by purging
i~ with the desorption effluent from the larger pore
sized adsorbent. Such operation minimizes the
amount of non-adsorbable purge gas reguired since
the two adsorber beds are configured in series.
Furthermore, such operation is feasible since the
desorption effluent from the larger pore adsorbent,
which comprises mono-methyl paraffins, is not
readily adsorbable on the calcium zeolite ~.
Another class of adsorbents, those having
eliptical pores with cross sectional pore dimensions
in the range of approximately 4.5 to 5.5 Angstroms,
are useful for separating isopentane from a
hydrocarbon feed containing a pentane fraction.
Preferred adsorbents for this application are ZSM-23
and ZSM-ll and are described in U.S. Pat. No.
4,076,872 and U.S. Pat. No. 3,709,979. The zeolite
ferrierite, as described in U.S. Pat. No. 4,016,425
and U.S. Pat. No. 4,251,499, is also a suitable
adsorbent for separating isopentane from a feed
containing a pentane fraction.
S~
-13-
In this aspect of the invention, the
feed, which contains iso and normal pentane, i
passed throuqh an adsorber bed containing the above
mentioned adsorbent wherein both isopentane and
normal pentane are adsorbed and isopentane i~
preferentially desorbed during continued adsorption
to provide a mass transfer zone having isopentane
concentrated at the leading edge thereof. The
adsorption can be continued so as to recover an
adsorption effluent and elute at least a portion of
the mass transfer zone, which comprises isopentane,
from the adsorber as an adsorption effluent.
Regeneration of the adsorbent can be accomplished by
desorption by any conventional means including
non-adsorbable purge, displacement purge,
temperature swing, and pressure swing.
The operating conditions of the adsorbers
are generally not critical to performance of the
process. However, it will usually be desirable to
operate the adsorbers in the temperature range of
40~C to 400C, preferably between 100C and 260C,
and in the pressure range of 3 ~ar to 40 bar,
preferably between 10 bar and 20 bar.
In addition to the adsorption aspects of
the invention, in many instances it will be
advantageous to incorporate an isomerization reactor
into the process to isomerize the adsorbed normal
and mono-methyl paraffins. The isomerization
catalyst suitable for use in the process of the
present invention includes all catalysts capable of
isomerizing normal and mono-methyl branched
2 ~ 2 ~
-14-
paraffins to more highly branched chain paraffin~.
One type can be any of the various molecular
sieve-based.catalyst compositions well known in the
art which exhibit selective and substantial
isomerization activity. As a general class, such
catalysts comprise the crystalline zeolitic
molecular sieves having apparent pore diameters
large enough to adsorb neopentane; SiO2/A1203
molar ratios of greater than 3; and less than 60,
preferably less than 2~, eguivalent percent alkali
metal cations wherein those A104 tetrahedra not
associated with alkali metal cations are either not
associated with any metal cation, or associated with
divalent or other polyvalent metal cations.
Exemplary of such zeolites are mordenite and
zeolite ~.
Because the feedstock may contain some
olefins and will undergo at least some cracking, the
zeolitic catalyst is preferably combined wi~h a
hydrogenation catalyst component, preferably a metal
of group VIII of the Periodic Classification of the
Elements (Periodic Table of the Elements, Handbook
of Chemistrv and PhYsics, 46th edition, The Chemical
Rubber Co., 1965-1966). The catalyst composition
can be used alone or can be combined with a porous
inorganic oxide diluent as a binder material.
Suitable catalysts of this type are disclosed in
detail in U.S. Pat. Nos. 3,236,761 and 3,236,762.
One such catalyst is prepared from a zeolite Y (U.S.
Pat. No. 3,130,007) having a SiO2/A1203 molar
ratîo of about 5 by reducing the sodium cation
content to less than about 15 equivalent percent by
2~2~3
-15-
ammonium cat;on e~c~ange, then introducing between
about 35 and 50 equivalent percent of rare earth
me~al cations by ion exchange and thereafter
calcining the zeolite to effect substantial
deammination. As a hydrogenation component,
platinum or palladium in an amount of about 0.1 to
1.0 weight percent, can be placed on the zeolite by
any conventional method.
Another suitable type of isomerization
catalyst which may be used in this process is a
composite of a metal from group VIII of the Periodic
Table with a solid support. The platinum group
metals, and platinum in particular, are preferred
for use as components of such a catalyst in this
process. Solid supports which are suitable for a
catalyst of this type include silica, alumina,
magnesia, zirconia, chromia, etc. A preferred
support or carrier material is alumina, It. is also
preferred that this type of catalyst incorporate a
halogen component into the catalyst composite to
impart additional acidity and activity to the
catalyst. This combined halogen may be either
fluorine, chlorine, iodine, bromine, or mixtures
thereof. Chlorine is the most preferred halogen for
purposes of the present invention and will be
typically combined with the carrier material
sufficient to result in a final composite that
21~2~3
-16-
contains preferably about G.1 to 5.0 total wei~ht
percent haloqen. In addition, small amount~ of
halogen, such as in the form of carbon
tetrachloride, may be continuously added to the
catalyst to offset any halogen loss. In some
instances, it may be benef icial to impregnate the
catalyst with an anhydrous Friedel-Crafts type metal
halide, such as aluminum chloride. Catalysts of
this type are disclosed in detail in U.S. Pat. Nos.
2,999,074, and 3,772,397.
In qeneral, the isomerization reaction can
be carried out in the vapor phase when either of the
two catalyst types is used, although complete
vaporization does not need to be maintained when the
haloqenated catalyst is used. The operating
temperature of the isomerization reactor is
~enerally within the range of 40C to 400~C and,
more specifically, in the range of 100C to 200C
for the halogenated catalyst and 200C to 400C for
the noble metal zeolite catalyst. The operating
pressure is typically within the range of 3 bar to
40 bar but is not critical to the isomerization
performance. However, it is generally desirable to
operate the isomerization reactor in the range of 10
bar to 20 bar in order to be compatible with the
adsorption section which preferentially operates in
the range of lO bar to 20 bar. The hydrocarbon
flowrate through the isomerization reactor is
generally maintained at a weight hourly space
2~2~ 3
-17-
velocity from 0.5 to 5.C hr 1, and, more
typically, from 0.5 to 3.0 hr 1. In order to
prevent catalyst coking a hydrogen partial pressure
in the range of 3 to 20 bar, and, more preferably,
in the range 6 to 14 bar should be maintained over
the isomerization catalyst.
When the adsorption process of the present
invention is practiced in conjunction with the
isomerization process of the present invention,
several configurations are possible. One
configuration is known as a reactor lead process and
is described in U.S. Pat. No. ~,210,771. In the
reactor lead process, the feed is passed through the
isomerization reactor prior to being passed through
the adsorbers. Desorption effluent from the
adsorbers is then combined with feed and passed
through the isomerization reactor. Another
configuration is known as the adsorber lead process
and is described in U.S. Pat. No. 4,709,116. In the
adsorber lead process, the feed is passed through
the adsorbers prior to being passed through the
isomerization reactor. At least a portion of the
effluent from the isomerization reactor is then
combined with feed and passed through the adsorbers.
A third configuration involves a split feed process
wherein a portion of the feed is initially passed
through the isomerization reactor, and another
portion is initially passed through the adsorbers.
2~2~3
-18-
A preferred combination according
to the present invention is to use one of the larger
pore, high silica adsorbents described suPra, e.g.,
silicalite or ZSM-5, with a halide activated
isomerization catalyst. Because of their high
silica-low alumina content, adsorbents of this type
possess a unique acid-resistant characteristic and
removal of the halides is not reguired. Another
preferred combination is to use one of the larger
pore, high silica adsorbents such as silicalite or
ZSM-5 in conjunction with a smaller pore calcium A
zeolite adsorbent as hereinbefore described with an
isomerization catalyst that is not halide activated.
Description of the Drawinq
FIG. l is a schematic of a process
according to ~he invention wherein an integrated
adsorption-isomerization process is used to upgrade
the octane of a hydrocarbon feed comprising a
pentane and hexanè paraffin fraction.
-19- 2~ iJ~
Example
For purposes of illustrating the invention,
the following description and example of an
integrated adsorption-isomerization process is
provided in conjunction with the drawing. The
fresh feed to the process has a Research Octane Num~er
of 64 and the following composition:
Component Mole Percent
iC4 0.2
nC4 4.0
iC5 18.4
nC5 28.9
CP (cyclopentane) 1.2
22DMB 0~4
23DMB 1.0
2MP 7
3MP 4 5
nC6 26.3
MCP (methylcyclopentane) 4.1
CH (cyclohexane) 3.0
BZ (benzene) 0.4
C7+ 0.2
-20- 2~2~
The fresh hydrocarbon reed is fed via Line 10 at a
rate of 280.7 kgmol/hr to Valve 216 where it mixes with the
reactor effluent condensate from Line 114, to form the
adsorber feed at a rate of 799.1 kgmol/hr. The adsorber
feed in Line 11 is heated by indirect heat exchange with
reactor effluent in Heat Exchanger 12 from which it is
passed to Furnac~ 14 where it is heated to approximately
260C for passage to the adsorption section via Line 16.
Adsorber feed from Line 16 and Furnace 14 are
directed partially to Line 20 by way of Pressure Control
Valve 22 and partially to Line 24 by means of Flow Rate
Control Valve 26. Each of the four adsorber vessels, 30,
32, 34 and 36, contains a dual bed of adsorbent with
30849 kg of calcium zeolite A adsorbent in the form of
1.6 mm (1/16 inch) in diameter cylindrical pellets in the
feed end bed of the adsorber vessel and 24092 kg of
silicalite adsorbent in the form of 1.6 mm diameter
cylindrical pellets in a second bed at the outer end of the
adsorber vessel. In a four vessel system, each of the beds
in the vessels cyclically undergoes the stages of:
A-l adsorption fill, wherein the vapor
in the bed void space consists principally of a
non-adsorbable purge gas, preferably hydrogen,
is forced from the bed void space by the
incoming hydrocarbon feed without
.
2~2~3
-21-
substantial intermixing thereof with
nonadsorbed feed fraction;
- A-2 adsorption, wherein the feed is
cocurrently passed through said bed and the
normal and mono-methyl paraffins including
isopentane in the feed are selectively
adsorbed while the nonadsorbed compone~ts
of the feedstock are removed from the bed
as adsorption effluent and during such
adsorption isopentane is preferentially
desorbed during continued adsorption to
provide a mass transfer zone ha~ing
isopentane concentrated at the leading edge
thereof, and eluting at least a portion of
the mass transfer zone, which comprises
isopentane, from the adsorber bed as
adsorption effluent;
D-l purging, wherein the adsorbent
which is loaded with essentially normals in
the calcium zeolite A section and
essentially hexane mono-methyl paraffins in
the silicalite section and containing in
the bed void space essentially a feedstock
mixture, is purged countercurrently, with
respect to the direction of A-2 adsorption
by passing through the bed a nonadsorbable
purge gas, preferably hydrogen, in the form
of a hydrogen-containing recycle stream
which comprises hydrogen and light
hydrocarbons, in sufficient quantity to
remove said void space feedstock vapors in
the bed effluent;
2~2~
-~2-
D-2 purge desorption, wherein the
selectively adsorbed feedstock normals and
mono-methyl paraffins are desorbed as part
of the desorption effluent by passing a
nonadsorbable purge gas, preferably
hydrogen in the form of a
hydrogen-containing recycle stream which
comprises hydrogen and light hydrocarbons,
through the bed countercurrently with
respect to the A-2 adsorption s~ep until
the major portion of the adsorbed
components have been desorbed and the bed
void space vapors consist principally of
the nonadsorbable purge gas.
~eferring again to the drawing, and the adsorpti~n
section in particular, the following description
details an operation wherein Bed 30 is undergoing
A-l adsorption-fill; Bed 32, A-2 adsorption; Bed 34,
D-l void space purging and Bed 36, D-2 purge
desorption. A portion of the adsorber feed from
Line 16 is directed via Line 24 through Manifold 40
and Valve 42 to Adsorption Bed 30 undergoing A-l
adsorption.
Bed 3D, at the time that feed passing
through Valve 42 enters, contains residùal purge gas
from the preceding desorption stage. The purge gas
is typically hydrogen-containing because of the
desire to maintain at least a minimum hydrogen
partial pressure in the isomerization reactor. This
purge gas is supplied to the adsorbers during
desorption as a purge gas recycle stream via Line
80. The rate of flow of the adsorber feed through
e~3
-23-
Line 24, Manifold 45 and Valve 42 is controlled such
that Bed 44 is often flushed of residual
hydrogen-containing purge gas adsorber stage time
period of, for instance, from about thirty seconds
to about two minutes.
During this first stage of adsorption in
Bed 30, the hydrogen-containing purge gas effluent
passes from the bed through Valve 50 into Manifold
s2. During the time period when the
hydrogen-containing purge gas is being flushed from
Bed 30, the remaining adsorber feed passes through
Valve 22 and Line 20, through Manifold 44 and Valve
46 to Bed 32.
The normal and mono-methyl paraffins are
adsorbed by Bed 32 which is undergoing A-2
adsorption and an adsorber effluent, i.e., the
nonadsorbed non-normals and adsorbed isopentane
emerges from the bed through Valve 54 and from there
is fed to Manifold 56. The adsorber effluent flows
through Line 60, Heat Exchanger 62 and Line 64 at a
rate of 389.7 kgmol~hr and is then further cooled to
38C in Heat Exchanger 67 and flows through Line 65
to Separator 268 where the liquid product is
withdrawn at a rate of 309.0 kgmol/hr and the
overhead vapors flow through Line 269 to be recycled
as purge gas.
The purge gas from Line 269 at a rate of
80.7 kgmol/hr is combined with the purge gas from
Line 69 and make-up purge gas from Line 120 at a
rate of 30.2 kgmol/hr and passed through Compressor
70 at a combined rate of 1498.2 kgmol/hr.
-~4- 2~
From Compressor 70, the hydrogen-containing
purge gas is passed through Line 72 and Heat
Exchanger 6~ and Heater 74, wherein it is heated to
a temperature of approximately 2~0~C and then passed
through Line 80 as the pur~e gas recycle ~tream.
The pressure of the adsorbers will typically be
within the range of from 13.6 to 21.8 bar, and
preferably will be in the range of from 16.3 to 20.4
bar.
The hydrogen-containing purge gas recycle
stream from Line 80 can be divided into two streams
by means of Flow Control Valves 82 and 84, and the
lesser stream passed through Line 86, Manifold 88
and Valve 58 countercurrently (with respect to the
previous adsorption stroke) through Bed 34. The
1 QW, controlled flow rate employed for the first
stage desorption is for the purpose of flushing
nonadsorbed hydrocarbons from the bed voids without
causing excessive desorption of the normals from the
adsorbent.
The effluent from Bed 34 passes through
Valve 48 and into Manifold 40 where it is recycled
through Valve 46 directly to Bed 32 undergoing A-2
adsorption. The major portion of the hydrogen
recycle stream from Line 80 is passed through
Control Valve 82, Line 90, to Manifold 52 where it
is mixed with the previously mentioned first stage
adsorption effluent from Valve 50 and then passes
through Valve 92 and Bed 36. During this stage,
selectively adsorbed normal and mono-methyl
paraffins are desorbed from the molecular sieve and
flushed from the bed. The desorption effluent from
-25-
Bed 36, compri~ing hydrogen and desorbed paraffin~,
passes through Valve 94 and Manifold 96 to Line 100,
from which .it is sent to the Isomerization Reactor
102 as reactor feed.
The foregoing description is for a single
adsorber stage time period of a total four-stage
cycle fsr the system. For the next adsorber sta~e
time period, appropriate valves are operated so that
Bed 30 begins A-2 adsorption, Bed 32 begins D-l
purging, Bed 34 begins ~-2 de~orption and Bed 36
begins A-l adsorption. Similarly, a new cycle
begins after each adsorber stage time period; at the
end of the four cycle time periods, all the beds
have gone through all stages of adsorption and
desorption.
The followins chart illustrates the
functioning of each of the four beds for adsorption
stage cycle times of one minute.
-
Time, Min. 0-1 1-2 2-3 3-4
-
Bed 30 A-l A-2 D-l D-2
Bed 32 A-2 D-l D-2 A-l
Bed 34 D-l D-2 A-l A-2
Bed 36 D-2 A 1 A-2 D-l
_
The isomerization process will result in
some hydrogen losses from the purge gas due to
hydrogenation of starting materials and cracked
residues. Hydrogen will also be lost due to
solubility in product, and possibly a vent from Line
69 (not shown) which can be controlled by suitable
2 ~ 2 ~
-26-
valve means. These losses require the addition of make-up
hydrogen. Make-up hydrogen can be supplied in impure form,
typically as an off-gas from catalytic reforming or stream
reforming of methane These hydrogen sources are suitably
pure for isomerization processes which typically have a vent
from the recycle stream. ~efinery streams of lesser purity
may also be satisfactory. The desorption effluent in Line
100 will comprise desorbed normal and mono-methyl paraffins,
e.g., n-pentane, n-hexane, 2-methylpentane, 3-methylpentane
and hydrogen and light hydrocarbons and other impurities
which comprise the purge gas used for desorption. This
effluent is reactor feed and is passed to Isomerization
Reactor 102.
Isomerization Reactor 102 contains a mordenite base
isomerization catalyst. The effluent from Reactor 102 flows
via line 104 through Heat Exchanger 12 and via line 108 to
Water Cooler 112 at a rate of 1905.7 kgmol/hr to Separator
68 where liquid is withdrawn through line 114 and combined
with fresh feed at Valve 216.
2~2~
-27-
The liquid product withdrawn from Separator 268
via line 214 after stabilization has a Research Octane
Number of 93.5 RON and the following composition:
Component Mole Percent
iC4 2.6
nC4 1.5
iC5 52.4
nC5 0.5
CP 1.4
22DMB 19.7
23DMB 11.0
2MP 0.7
3MP 1.5
nC6 0.5
MCP 4.4
CH 3.1
BZ 0.4
C7+ 0.3
