Note: Descriptions are shown in the official language in which they were submitted.
2 ~ 7 ~
ENERGY EFF I C I ENT ARO~T I C EXTRACT I ON PROCES S
_ELD OF THE INVENTION
The invention relates to an improved, more .
energy 0fficient process for the separation of
aromatic and nonaromatic hydrocarbons ~rom a mixed
hydrocarbon feed, and more particularly, to the
separation of aromatic and nonaromatic hydrocarbons
in high yields from a mixed aromatic, naphthenic and
paraffinic hydrocarbon feed. The instant process
significantly decreases the energy reguirement~
necessary for the separation of aromatic and
nonaroma~ic hydrocarbons and especially well suited
for separation of lube oil fractions.
BACKGROUND
-
The separation of aromatic and nonaromatic
hydrocarbons (generally referred to as
dearomatization) from mixed hydrocarbon feeds has
long been recognized as necessary and advantageous
for a number of varied reasons. For example, when a
BTX fraction (benzene, toluene and xylene) is the
aromatic fraction it may be used as a raw material
in the manufacture of petrochemicals, or as an
additive for gasoline to increase its octane
rating. Further, the nonaromatic fraction derived
from these mixed feeds have varied uses as fuels,
solvents and the like and, therefore, are also
highly desirable. Such uses for the aromatic and
nonaromatic fractions have resulted in the
development of numerous dearomatization processes.
Of particular interest and difficulty is
the separation of the complex components present in
D~ 1 3 , 8 1 7
- 2 - ~3~2~7~
lube oils, wherein the removal of aromatic-type
hydrocarbons is necessary to improve the viscosity
index, thermal and oxidation stability, and color of
the lube oils. The presence of aromatic-type
hydrocarbons in lube oils affects the quality of
these oils due to the low viscosity index, poor
thermal and oxidation stability, high carbon
residue, and poor color of such aromatic-type
hydrocarbons. The aromatic-type hydrocarbons
present in lube oils differ significantly from the
BTX fraction found in light hydrocarbon mixtures
used in the production of gasoline and, as a result,
present vastly different separation problems.
Various processes have been suggested for
the separation of the aromatic and nonaromatic
hydrocarbons of a mixed feed wherein the aromatic is
a BTX fraction. Typical of these processes is a
process employing an extraction column for
separation of a BTX fraction wherein a selective
solvent, BTX and a reflux stream is introduced to a
two step distillation column. BTX is then distilled
to remove water and entrained solvent. Similarly, a
process has been suggested wherein two distillation
columns are employed with the BTX fraction and water
being distilled in the second column. In addition,
a process using two distillation columns wherein the
second column is employed to distill the BTX
fraction and other components, has been suggested.
One goal o~ the prior art has ~elated to
processes developing a dearomatization process which
lowers the cost of dearomatization of the mixed
hydrocarbon feed. This reduction in cost for
D-13,817
- - 3 ~ ~3~257~
dearomatization can be achieved by improving the
selectivity of the selective solvent and by
modification of the separation process scheme. U.S.
Patent 3,985,644 mentions one such method for
modifying the process scheme and reducing
dearomatization costs, i.e., by reducing the use of
energy-intensive steps, e.g., distillation.
The dearomatization of lube oils is of
particular interest. Dearomatizecl lubricating oils
are, generally speaking, naphthenic- and or
paraffinic-type viscous materials having a low rate
of viscosity change with change in temperature,
i.e., relatively high viscosity index, a high degree
of thermal and oxidation stability, low
carbon-forming tendency, good color, and high flash
points. Lubricating oil feedstocks are generally
recovered as distillates or bottoms from the vacuum
distillation o~ crude oils. A crude lube oil
fraction contains many differenk chemical
components, e.g., paraffins, naphthenes, aromatics,
and the like. In order to obtain refined
lubricating oils of relatively good quality and high
viscosity index, the practice has been to remove
components, such as aromatic and polyaromatic
compounds, which tend to lower the viscosity index
of the lube oil. The removal of these aromatic
components has heretofore been carried out by
processes as above-described and processes such as
disclosed in U.S. Patent Nos. 2,079,885-; 2,342,205;
3,600,302; 2,773,005; 3,291,728; 3,788,980; and
3,883,420.
D-13,817
-- 4 --
In EPO patent application publication No.
43,267 published January 6, 1982, a solvent extractlon~
solvent decantation process is disclosed wherein solvent
purificatlon with mixed hydrocarbon feed or ra~finate
can be employed.
In EPO patent application publication No.
43,685 published January 13, 1982, there is disclosed a
process for the separation of aromatic and ~onaromatic
containing feedstocks by use of a unlque extraction-
decantation process wherein the extraction solvent is
preferably a low molecular weight polyalkylene glycol.
The instant process provides for an improved process by
use of improved mixed extraction solvents.
~ he proce~ of thi~ invention i~ to be
distinguished from the above described proce~e~ in that
the in~tant procQss provide~ an energy balanced
extraction-separation proce~s that i~ more economically
advantageou~, i.e., energy efficient than the above-
described processes.
SUM~E~OF T~E INVEN~ION
In accordance with one aspect of th2 present
invention, there is provided a proce s for the
dearomatization of a mixed hydrocarbon feed (containing
aromatic and nonaromatic components~ with low energy
consumption in a continuous solvent extraction-solvent
separatio~ process, which ~omprises the following
steps:
(a) contacting the hydrocarbon feed with an
aro~atic extraction solv~nt in an extraction zone, at a
te~p~rature o~ at least about 15~C to provide a
solvent phase containing aro~atic hydrscarbons and a
raffinate phase containing nonaromatic hydrocarbons;
D 13,817
.~3
- 5 - ~31~
~ b) cooling the solvent and raffinate
phases;
(c) introducing said cooled solvent
phase to a separation zone and introducing therewith
about 0.1 to about 25.0 percent by weight of an
anti-solvent for said aromatic hydrocarbons in said
solYent phase to provide an extract phase containing
aromatic hydrocarbons and a solvent-rich phase
containing mixed extraction solvent and anti~solvent;
(d~ introducing said cooled raffinate
phase ~o a separation zone and introducing therewith
about 0.5 to about 75.0 percent by weight of an
anti-solvent for said aromatic selective solvent in
said raffinate phase to provide a raffinate phase
and a sol~ent/anti-solvent phase;
(e) removing anti-solvent from said
solvent-rich phases of steps (c) and (d) and
recycling said solvent-rich phase to the extraction
zone of step (a), and
(f) recovering the extract phase of
step (c) and the raffinate of step (d).
Small amounts of entrained and/or or
dissolved solvent may be removed from the
solventtanti-solvent containing raffinate and
aromatic extract phases by means of water wash
processes.
BRIEF DESCRIPTION OF THE DRAWING
Figure 1 is a schematic flow ~iagram of the
process of the invention.
Figure 2 is a schematic flow diagram of the
process of the invention wherein steam export is
shown.
D-13,B17
- 6 - ~ ~12~7~
ESCRIPTION OF THE INVENTION
There has historically been and continues
to be an industrial need for an ener~y efficient
process for the separation of arornatic and
nonaromatic hydrocarbons present in mixed
hydrocarbon feeds. Naphthas, heating oils, light
oils, cracked gasolines, dripolenes, lubricating
oils (light distillates to heavy distillates)
kerosene and the like, can contain up to 90 percent
by weight aromatic-type hydrocarbons, e.g., BTX or
polyaromatics. The separation of aromatic and
nonaromatic hydrocarbons is of particular interest
in the dearomatization of crude lube oils. The
components which make up these hydrocarbon eed
streams are well known in the art and they will not
be extensively discussed herein except to note that
the mixed hydrocarbon feed employed herein may be
any petroleum of the common distillation fractions
containing one or rnore aromatics components
including: naphthas (virgin or cracked); kerosene;
gasoline; heating oils; lubricating oils, (light
distillates heavy distillates, bright stock and
residual oils); jet fuels; and recycle oils.
Preferably, the feed stream is a lube oil fraction
such as a light distillates to heavy dis~illate,
bright stock, etc., which have boiling points
between about 400F and about 1100~F.
The aromatic hydrocarbons present in heavy
hydrocarbon feeds, e.g., lubricating oils, generally
include: alkylbenzenes, indanes, tetralins,
indenes, naphthalenes, fluorenes, acenaphthalenes,
biphenyls, phenanltrenes, anthracenes,
D-13,817
- 7 - ~`3
diacenaphthalenes, pyrenes, chripenes,
diaceanthrancenes, benzpyrenes and other various
aromatic feed components.
The instant process provides for
significantly improved processing of feedstocks
containing aromatic and nonaromatic components. Tne
instant process provides for an enerqy efficient
separation of aromatic and non-aromatic feed
components. The process makes use of special
extraction solvents and of separation and
decantation steps for both the aromatic ~extract)
and nonaromatic (raffinate) components of the mixed
hydrocarbon feed. The instant process may be
employed for treating different feedstocks, e.g.,
from light paraffin distillate to bright stock,
without significantly increasing the energy
requirements of the process. l'his results in
savings, in terms of the energy consumpti~n of the
process, of up to 50 percent or more, based on the
ener~y required in furfural or similar processes
presently employed in the field of such
separations. For example, furfural refining is
disclosed in "Hydrocarbon Proce~sing", p. 188,
September 1978, to involve economics wherein
treatment of 7000 bbl/day of a light oil feed would
re~uire over 73 million BTU/hour. The instant
process would require less than 40 million BTU/hr.
to treat such a feed. This signi f icant energy
savings is commercially important.
The solvents and cosolvents employed in the
instant process may be any of water-miscible organic
liquids (at process temperatures) having a boiling
D-13,817
~3~2~7~
point and decomposition temper~ture higher than the
extraction temperature and having selectivity for
aromatic compounds. The term "water-miscible" describes
those 501vent5 and cosolvents which are completely
miscible with water over a wide range o~ temparatures
and which have a high partial miscibility with water at
roo~ temperature, since the latter are usually
co~pletely miscible at process temperatures.
The aromatic ~xtraction solvent employed in
the process of the invention comprises a polyalXylene
glycol and a cJlycol ether. The polyalkylene glycol
component ("solvent") is one of the formula:
H0-[CHR1~(cR2R3)n~0]m-H
wherein n is a integer ~rom 1 to 5 and is preferably tne
integer 1 or 2; m is an integer haviny a value of 1 or
greater, preferably between about 2 to about 20 and most
pre~erably between about 3 and about 8; and wherein R1,
R~ and R3 may be hydrogen, alkyl, aryl, aralkyl,
alkylaryl and mixtures thereo~ and are preferably
hydrogen and alkyl having between 1 and about 10 carbon
atoms and most preferably are hydrogen. Examples of the
polyalkylena glycol solvents e~ployable herein are
d~ethylene glycol, triethylene glycol, tetraethylene
glycol, 1,3-butylenQ glyrol, 1,2-butylene glycol, 1,5-
pentaethylene glycol, and mixtures thereo~ and the
like. Preferred solvents are diethylene glycol,
D 13,817
~,
~ ~312~7~
g
triethylene glycol, tetraethylene g:Lycol, or mixtures
thereof with tetraethylene glycol being most preferred.
The glycol ether component ~"cosolvent") is one of the
formula:
R40- [CHRs- (CHR6~)-xo]~~ R7
wherein R4, R5, R6 and R7 may be hydrogen alkyl, aryl,
aralkyl, alkylaryl and mixtures thereof with the proviso
that R4 or R7 are not both hydrogen. The value of x is
an integer from 1 to 5, preferably 1 or 2 and y is an
integer from 2 to 10 and is preferably from 2 to 7, and
most preferably from 2 to 5. R4, R5, R6 and R7 are
pre~erably selected Erom the group consisting of
hydrogen and alkyl having 1 to about 10 carbons with the
proviso that R~ and R7 may not both be hydrogen and most
preferably R4 is alkyl having from 1 to 5 carbons and R5
R6 and R7 are hydrogen. The mixture(s) of solvent and
cosolvent is selected such that at least one solvent and
one cosolvent are provided to form the mixed extraction
solvent. The cosolvent generally comprises between zero
(0) and about 99 percent of the mixed extraction
solvent, preferably between about 0.5 and
D 13,817
......
- 10~ 73l
about 80 percent and more preferably between about
10 and about 60 percent by weight based on the total
weight of the mixed extraction solvent.
~ n effective amount of an anti-solvent may
be employed to modify the capacity and/or
selectivity of the solvent in the individual
separation steps of the two phases (extract and
raffinate) obtained from extraction zone and such
may be most any compound that tends to decrease the
solubility of the aromatic and non-aromatic
hydrocarbons in the extraction solvent. Water is
the preferred anti-solvent. The use of water as an
anti-solvent in the extraction step has been
observed to provide added process versatility by
correlating the selected feed with the extraction
solvent and the effective amount of anti-solvent
present in the extraction zone. Other suitable
anti-solvents are believed to include ethylene
glycol, glycerine, low molecular weiyht alcohols and
the like. llhe use of low molecular weight alcohols,
even at the low concentration of antisolvent
employed, is not generally preferred owing to the
art recognized problems associated with the use of
alcohols. The effective concentration of the
anti-solvent as determined in the extract separation
zone is that amount which effectively decreases the
solubility of the aromatic hydrocarbon in the
extraction solvent as determined by the amount of
aromatic hydrocarbon in the extraction solvent
leaving the separation zone, The concentration of
aromatic hydrocarbons in the extraction solvent
leaving the extract ssparation zone or the amount of
D-13,817
11- ~3~2~
solvent in the raffinate leaving the raffinate
separation zone is preferably less than 3 percent by
weight, based on the weight of the extraction
solvent in the aromatic extract or the weight of
solvent in the raffinate and is preferably
preferably less than 2 percent by ~weight and more
preferably less ~han 1 percent by weight. The
anti-solvent employed in the insta:nt process
promotes the formation of two phas~es to a degree
greater than that obtainable by simple cooling of
the phases obtained by extraction. The
afsrementioned cooling and addition of anti-solvent
results in the formation of an aromatic-rich extract
phase, a raffinate phase and two solvent-rich
phases. Generally the concentration of the
anti-solvent present in the extract separation zone
is in the range of from about 0.5 to 25.0 percent by
weight or higher, based on the total weight of the
aromatic-rich solvent phase, with the range from
about 0.5 to about 15.0 percent being preferred and
the range rom about 3.0 to about 10.0 being mo6t
preferred. The concentration of the anti-solvent
present in the raffinate separation zone is in the
range of from about 0.5 and about 75 percent by
weight, preferably from about 1.0 to about 50
percen~ by weight and more preferably from about 1.0
to about 40 percent by weight, based on the total
weight of the raffinate and anti-solvent in the
raffinate separation zone. Some portion of the
anti-solvent present in the separation zones
(extract and raffinate) may be provided by
anti-solvent which may be present in the solvent
D-13,817
- 12 - ~ 312~
phase employed in the extraction zone as a result of
amounts present in the solvent due to the recycle of
the extraction solvent and anti-solvent. As
hereinbefore noted, the presence of up to about 10
percent by weight (based on the total weight of the
recycled extraction solvent to the extraction zone~
may he advantageous in correlating the ielectivity
and/or capacity of the extraction solvent to changes
in the feedstock composition~ The actual
concentration of the anti-solvent in the separation
zones will depend at least in part on the nature of
the hydrocarbon feed, nature and relative amount of
aromatics present in the hydrocarbon feed, the
extraction solvent employed, the anti-solvent
employed and the like. The aforementioned
concentrations designate the total anti-solvent,
e.g. water, present in the separation zone
irrespective of its source. Anti-solvent is
preferably added to the raffinate and aromatic-rich
solvent phases prior to the respective separation
zone so as to provide for improved separation in the
respective separation zones. In addition, cooling
of the raffinate and separation of some extraction
solvent prior to the addition of anti-solvent may be
advantageous, e.g., more efficient, in some
instances in providing for the separation of solvent
from the raffinate.
Generally, to accomplish the extraction,the
ratio of the extraction solvent to hydrocarbon feed
in the extractor zone is in the range from about 2
to about 20 parts by volume of solvent to one part
by volume of feed, the ratio from about 2 to 1 to
D-13,817
-' ' ' .
j ~ ~ . j,
- 13 ~ ~3~2 5~ ~
about 15 to 1 being preferred and the ratio from
about 4 to 1 to about 10 to 1 being the most
preferred. The broad ranqe for the ratio of the
solvent to hydrocarbon may be expanded upon
depending on the solvent, co-solvent, if any, weight
percent of solvent to cosolvent, the weight percent
of anti-solvent in the mixed extraction solvent and
the like. The optimum solvent to feed ratio also
depends upon whether high recovery (yield) or high
purity (quality) is desired, although the instant
process will generally result in both high recovery
and high purity.
In one embodiment of the instant invention
it has been observed that by using a solvent ancl a
cosolvent as a "mixed extraction solvent" that the
process has a high selectivity and capacity for
aromatic hydrocarbons and, further, provides a
process wherein the heat duty for the process (the
energy consumption of the process) is proportional
to the rate at which the feed is introduced and not
strictly dependent on the ratio of the solven~ to
feed. Thus, the instant process provides for the
use of similar energy requirements for vastly
different feedstocks. This important feature is not
found in processes wherein furfural, N-methyl-2-
pyrrolidone and phenol have been employed as the
extraction solvent in processes not according to the
instant invention.
The instant process is further
characterized in that the pressure at the top of the
extraction zone is typically less than about 150
psig and often less than about 100 psig. This is
D-13,817
7 ~
highly advantageous in terms of ease of operation
and the capi.tal expenditure required for carrying
out the separation process. The actual pressures in
the extraction zone may be higher or lower depending
on the particular hydrocarbon feed treated, the
extraction solvent employed, the ~selected
antisol~en~ and its concentration, and the selected
temperature at which the extraction is carried out.
The pressure employed in the separation zones is
generally that pressure which is eequired to cause
the aromatic-rich solvent phase or raffinate phase
to pass through the separation zone, although higher
pressures may be employed if desired. Generally a
small pressure drop (pressure gradient) is observed
across the separation zones.
The temperature of the extraction zone is
generally at least about 150C and is generally in
the range of from about 150C to about 275C,
preferably in the range of from about 170C to about
~50C and most preferably from about 200C to about
240OC. The temperature in the extraction zone is
not constant throughout and there will generally be
a temperature gradient up to about 30OC or more,
typically from 15C to 30C, as between the
temperature of the mixed extraction solvent
introduced to the extraction zone and the
temperature of the solvent phase leaving the
extraction zone. The separation zones are generally
maintained at temperatures in the range of from
about 50C to about 200C below the temperature of
the extraction zone such that the temperature is
preferably in the range of from about 25C to about
D-13/817
- 15 - ~3~
150~C, more preferably about 25C to about 100C and
most preferably from about 25C to about 70OC. The
temperature employed in the separation zones
depends, in part, upon solubility of the aromatic
hydrocarbon and raffinate in the mixed extraction
solvent, the amount of anti~solvent present in the
separation zones and the viscosi~ 0f the extraction
solvent at the temperature(s) employe~ in the
separation zones.
The apparatus employed in the ;nstant
process in the extraction zone, separation zones and
otherwise are of conventional design. ~or example,
an extraction column of the multistage reciprocating
type containing a plurality of perforated plates
centrally mounted on a vertical shaft driven hy a
motor in an oscillatory manner can be used as may
columns containing pumps with settling zones or
sieve trays with upcomers or downcomers,
(Counter-current flow is generally utilized in the
extraction column.) The separations in ~he
separation zones can be conducted in a tank with no
internal elements but preferably the tank contains
coalescing elements or baffles to aid in the
separation. The preferred separation zone eomprises
a coalescer with a porous media having a depth~type
coalescing element (fibrous bed coalescer element).
It is understood that the "separation zone" is a
zone wherein phase separation is facilitated and
wherein anti-solvent is present. As noted
hereinabove, the anti-solvent is preferably added
prior to the separation zones after the solvent
phase or reaffinate phase exit the extrac~ion zone
and has been at least partially cooled.
D-13,817
- 16 ~ 2~7~
Heat exchangers, reservoirs, and solvent
regenerators, if necessary, are al50 of conventional
design as well as are the various extractors and
decanters used in the various embodiments
h~reinafter described. The extrac~ors employed are
preerably multi-stage counter-current extractors,
but can be any of the well-known t:ypes, as
a~orementioned.
The instant process generally provides for
an overall recovery of the aromatic hydrocarbon of
from about 70 to about 95 percent or better based
upon the weight of aromatic in the original
hydrocarbon feed and usually provides for similar
recoveries for the nonaromatic hydrocarbons.
The instant process may be carried out in
many process schemes accordin~ to the aforementioned
disclosures and consistent with heretofore generally
accepted process design principles. Figure 1
depicts a flow diagram of the instant process where
water is the anti-solvent and s~eam is employed
internally in the process. Figure 1 does not employ
steam export. A process according to this invention
involving steam export is depicted in Figure 2.
Referring to Figure 1
The mixed hydrocarbon feed is introduced at
lQ to pump 14 and split to lines 12A and 12B. The
feed passes through line 12A and heat exchanger 16
and the feed in line 12B heat exchanges with
raffinate in heat exchanger 18. At this point the
feed in lines 12A and 12B is combined and heated
with raffinate from line 26 in hea~ exchanger 20
before introduction to extraction column 24.
D-13,817 .,
- 17 - ~3~2~7 ~
Exchangers 16 and 18 are optional when the
temperature in extractors 38 and 39 is close to tha~
of the feed in line 12. An extraction solvenk,
preferably having a temperature in ~he range of from
about 150C to about 275C, most preferably about
200C to about 240C is introduced near the top of
extraction column 24 via line 57 and percolates down
column 24 removing aromatics from the hydrocarbon
feed forming a raffinate and an aromatic-rich
solvent phase. The raffinate, containing primarily
non-aromatics and aromatic selective solvent (e.g.,
about 2 to about 10 wt.~) exits the top of the
column 24 via line 26 and heat exchanger 20 where it
is cooled with the incoming hydrocarbon feed in line
12. The raffinate phase is further cooled in cooler
33, if necessary, to promote the formation of a
raffinate phase and a solvent phase. Fecycled
aromatic selective solvent and antisolvent from t:he
raffinate (extractor 3~) and extract washings
(extractor 38), when water is the selected
anti-solvent, are combined and introduced via line
44A to decantation tank (zone) 34A. The
anti-solvent is preferably added and thoroughly
mixed with the raffinate phase prior to decantation
tank (zone) 34A to further promote the formation of
two phases, e.g., at 46A of the drawing, in the
decantation zone, although ~he anti-solvent may be
add~d directly to decantation tank ~zone) 34A if
desired. It is understood that decantation
(separation) zones 34 and 34A may actually comprise
one or more separation stages which may involve one
or more separation apparatus and which may be
D-13,817
- 18 - ~ ~3~2~
carried in a countercurrent or other suitable
manner. It may be advantageous to cool and decant
solvent from the incoming raffinate prior to the
addition of anti-solvent and the further separation
of solvent. Water from the raffinate washings
(extractor 39) may be introduced at 46A if desired.
This will avoid even minimal contamination of the
raffinate phase in line 27 with the extract phase
from line 36.
The anti-solvent in the solvent/anti-solvent
mixture of line 44A reduces the solubility of the
aromatic selective solvent in the raffinate phase to
a degree not obtainable by simple cooling the
raffinate phase. The antisolvent is presen~ in
decantation tank (zone) 34A at a concentration from
about 1.0~ to about 75.0% by weight and most
preferably from about 1.0% to about 50.0% hy weight,
as above discussed. The presence of the
anti-solvent decreases the solubility of the solvent
in the raffinate phase such that typically less than
about 2% by weight solvent and often less than about
1% by weight solvent leaves decantation tank (zone)
34A via line 27. The raffinate then passes to
extractor 39 where it is contacted with water to
recover the remaining aromatic selective solvent
present in the raffinate so as to form a water phase
(containing aromatic selective solvent) and a final
raffinate product. A second water extraction takes
place in extractor 3B wherein the aromatic-riah
extract from decantation zone 34, discussed
hereinafter, forms a water-phase (containing
aromatic selective solvent~ and a final aromatic
D-13~817
- 19 ` ~L3~2:~7~
product. Alternatively, steam in line 54 (from
distilla~ion column 52) can be used in exchanger 22
to heat the raffinate phase of line 27 and/or the
aromatic-rich extract phase of line 36 prior to the
water-wash in extractors 39 and 38 (phantom lines in
Figure). This could be especially advantageous when
heavy feedstocks are being treated, e.g., bright
stock, ~o reduce the viscosity of the raffinate and
extract phases. The water-phases from extractors 38
and 39 contain primarily water and small amounts of
aromatic selective solvent that was dissolved or
entrained in the aromatic-rich extract and
raffinate. The combined wat~r-phases are recycled
to decantation tank (zone) 34A via line 44A, as
needed, if water is the selected anti-solvent and
from decantation tank 34A via line 44 to 46 prior to
introduction of the the aromatic-rich extract phase
to decantation zone 34.
It should be pointed out that the terms
"phase" and "product" are named after their main
components, which is present in the phase in an
amount of at least ~0% by weight and in many cases
in an amount of 90% by weight or higher. The
aromatic-rich solvent phase, containing primarily
aromatic selective solvent and aromatic
hydrocarbons, leaves the bottom of extraction column
24 via line 28 and heat exchangers 29 and 30 where
it is cooled with aromatic selective solvent in
lines 57 and 48, respectively. The aromatic-rich
solvent phase is further cooled to promote two phase
formation, if necessary, in cooler 32. Recycled
aromatic selective solvent and anti-solvent, when
D-13,817
- 20 - ~3~2~7~
water is the selected anti-solvent, are introduced
via line 44 to decantation tank (zone) 34. Thus,
the solvent contained in line 44 is returned to the
process. The anti-solvent is preferably added to
the aromatic-rich solvent phase prior to decantation
tank (zone~ 34 to further promote the formation of
two phases, e.g., a~ 46 of the drawing, in the
decantation zone, although anti-solvent may be added
directly to decantation tank ~zone) 34 if desired.
The anti-solvent in the solvent/anti-solvent mixture
of line 44 reduces the solubility of the aromatic
hydrocarbon in the aromatic selective solvent to a
degree not obtainable by simple cooling of the
aromatic-rich solvent phase. The anti-solvent
present in the decantation tank ~zone) 34 is
typically at a concentration of from about 0.5% to
about 25.0~ by weight, based on the weight of
aromatics and solvent in decantation tank (zone) 34,
preferably from about 0.5% to 15.0~ by weight and
most preferably from about 5% to about 10.0% by
weight. The concentration of the anti-solvent in
decantation tank (zone) 34A will be up to about 75
weight percent based on the weight of the raffinate
and anti-solvent. The presence of the anti-solvent
decreases the solubility of the aromatics and
raffinate in the extraction solvent such that
typically less than about 2 weight percent aromatics
in the ~olvent or solvent in the raffinate is
present and preferably less than about 1 weight
percent, leaves decantation tanks 34 and 34A,
respectively, via lines 48 and 27, respectively.
D-13,817
- 21 ~ 2~ ~ ~
.
The aromatic-rich extract phase of
decantation tank Szone) 34 exits via line-36 to
water-extraction column 38 where it is contacted
with water (preferably water derived from the
removal of water from the solvent~anti-solvent
mixture when water is the anti-solventj from the
solvent phase of decantation tank (zone) 34. This
extraction with water removes entrained and/or
dissolved solvent from the aromatic-rich extract
phase. Similarly, the raffinate phase of
decantation tank 34A exits via line 27 to
water-extraction column 39 where it is contacted
with water to remove entrained and dissolved solvent
from the raffinate phase.
The solvent phase of decantation tank
(zone) 34 passes via line 48 through heat exchanger
30 wherein it heat exchanges with hot aromatic-rich
solvent of line 28 prior to introduction to
distillation column (zone) 52. The solvent/water
phase from decantation tank 34A will typically
comprise up to about 90 weight percent water (when
water is the anti-solvent) and is advantageously
employed by introduction at 46 to the aromatic rich
solvent prior to or concurrent to its introduction
to decantation zone 34. The use of a distilla~ion
zone in the instant embodiment at 52 is not intended
to he limiting, since any means for decreasing the
concentration of the anti-solvent in the
aromatic-selective solvent may be employed. The use
of a distillation zone is preferred when the
anti-solvent is water, since the steam generated
therein may be advantageously and economically
D-13,817
~ 3 ~
employed in this and/or other processes (steam
export is not shown in FIG. 1).
When water is the selected anti-solvent the
solvent phase in line 4B is introduced ~o
distillation zone 52 wherein water is distilled,
preferably under pressure when generated steam is to
be exported to other processes, and removed as steam
via line 54. Steam in line 54 may be heat exchanged
at 22 with the mixed hydrocarbon feed, raffinate
and/or extract, after which the steam may be
condensed by cooler 62, and the water condensate may
be employed in extractors 38 and 39. Alternatively,
the steam leaving heat exchanger 22 may be
advantageously employed in this or other processes
(not shown). The aromatic selective solvent leaves
distillation zone 52 via line 57, it is heated first
in exchanger 29 with hot solvent from line 28
leaving extraction column 24, and finally it is
heated to extraction temperature in heater 31 before
going into extraction column 24.
As above discussed alternative schemes may
be substituted for that above--described for the
removal o~ the anti-solvent, depending on the
selection of the anti-solvent. Further, alternative
schemes employing steam export to other processes
may be employed as applicable to the various designs
of manufacturing facilities. The above described
scheme is particularly advantageous in terms of the
reduction in energy reguired to carry out the
instant process. For example, the above-described
extraction-separation process results in a reduction
in energy requirements for the dearomatization, as
D-13,817
. - 23 - ~3 ~2~7~
compared to conventional dearomatization processes,
by as much as 50 percent to about 80 percent
depending on the feed being subjected to
dearomatization by use of the process hereunder.
The total anti-solvent, e.g. water, in the
system can be easily determined because the amount
of water introduced at 46 and 46A to decantation
tanks 34 and 34A can be monitored. Allowances will
necessarily be made for water losses through
leakage, upsets, steam export, if any, and the like
so as to maintain the amount of anti-solvent present
in decantation tanks 34 and 34A at the desired
concentration, as above discussed.
The process described in FigurP l was
evaluated using a mixed extraction solvent of
tetraethylene glycol and methoxytriglycol. The
chemical analyses of the contents at selected points
in the process, as described in reference to Figure
l, are set forth in Tahle I below. Table I shows a
material balance for one embodiment of the process
at 60F using a bright stock feed.
Referring to Figure 2:
Figure 2 depicts a process according
to the instant invention wherein Steam export is
included as a feature of the process embodiment.
The mixed hydrocarbo~ feed is introduced at 80 and
passes through heat exchanger 82 where the incoming
feed and raffinate in line 90 are heat exchanged to
heat the feed and cool the raffinate. The feed is
introduced via line 84 to extraction column 86 where
the feed is contacted wi~h extraction solvent
introduced via line 144 to extraction column 86.
D-13,817
L 2 ~ 7 ~
The raffinate product exits extraction column 86 via
line 90, heat exchanger 82 and passes via line 92 to
heat exchanger 94 where the raffinate is cooled,
e.g., the raffinate may be heat exchanged with cold
water. The cooled raffinate is then introduced to a
raffinate decantation zone. The raffinate is
introduc~d to the decantation zone with an
anti-solvent effective in decreasing the solubility
of the non-aromatic hydrocarbons of the raffinate in
the extraction solvent. The anti-solvent,
preferably water, is typically added at a point
prior to the decantation zone, e.g., at 98, to allow
for thorough mixing with the raffinate to further
promote the formation of two phases in decantation
zone 100. A raffinate phase containing a higher
concentration of non-~aromatic hydrocarbons and a
lower concentration of extraction solvent than
present in line 92 exits decantation zone 100 via
line 102. The anti-solvent/solven~ phase from
decantation tank 100 may be introduced via line 104
at 106 to decantation zone 108. The aromatic-rich
extract phase leaves extraction column 86 via line
88 and via heat exchanger 112 and is cooled at 89.
Anti-solvent is added to the cooled aromatic-rich
solvent phase, preerably at 106 prior to
introduction to decantation zone 108, to promote the
formation of two phases. As above noted, the
anti-solvent introduced at 106 is preferably at
least in part derived from the anti-solvent/solvent
phase from decantation zone 100. Two phases are
formed in decantation zone 108. The aromatic-rich
phase exits via line 142 to water-extraction column
D-13,817
- 25 ~ 2 ~ 7 ~
134. Similarly, raffinate in line 102 is introduced
to water-extraction column 132. The raffinate and
aromatic rich extract are contacted with water to
remove entrained and/or dissolved solvent. The
raffinate is collected via line 146 and the extract
is collected via line 158. Wash water.from water
extraction columns 132 and 134 may be 0mployed via
lines 140 and 136, respectively, combined at 140 as
the anti-solvent and introduced via 96 to 98. The
solvent-rich phase from decantation ~one 108 exits
via line 110 and is typically heat exchanged at hea~
exchanger 112 with the aromatic-rich phase in line
88. The solvent~rich phase is then introduced via
line 114 to a distillation zone 116 where water is
distilled under pressure and exits via line 118 to
120 where steam is produced from condensate
introduced via line 128 for export to other
processes (not shown) via line 130. A rehoiler 122
may be employed. The solvent-rich phase is then
recycled via line 144 to extraction column 86.
D-13,817
- 26 - ~L3~7~
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