Process for the separation of hydrocarbons from a mixed feedstock

In a steam distillation process for the recovery of aromatic hydrocarbons wherein there is (i) a primary flash zone at the top of the distillation zone in which rich solvent is flashed and/or (ii) provision for the removal of side cut distillate vapors from about the middle of the distillation zone, the improvement comprising (a) heat exchanging flashed rich solvent vapors or side-cut distillate vapors with stripping water to provide stripping water vapors and stripping water at at least about the boiling point of water; (b) passing the stripping water vapors from step (a) to a steam ejector; (c) passing the stripping water from step (a) to a motive steam generator wherein the stripping water is vaporized under pressure; (d) passing the stripping water vapors from step (c) to the steam ejector referred to in step (b); and (e) passing the stripping water vapors, introduced into the steam ejector in accordance with steps (b) and (d), to the lower half of the distillation zone.

TECHNICAL FIELD 
This invention relates to a steam distillation process for the recovery of 
hydrocarbons from a mixed feedstock. 
BACKGROUND 
The benzene-toluene-C.sub.8 aromatic fraction (known and hereinafter 
referred to as BTX) is now well established as a premier raw material in 
the manufacture of petrochemicals and as a desirable component in boosting 
octane ratings in gasoline. Many processes have been proposed for the 
separation of BTX, e.g., the process proposed in U.S. Pat. No. 3,714,033, 
which is incorporated by reference herein. 
There is an industrial need for BTX, which is available in high proportion, 
e.g., greater than 30 percent by weight, in a wide variety of hydrocarbon 
feedstocks such as reformed gasolines; coke oven light oils; and cracked 
gasolines. These feedstocks also contain both aliphatic and cycloaliphatic 
hydrocarbons. Since the individual hydrocarbon compounds which make up 
these feedstocks are well known, they will not be discussed extensively; 
however, it can be pointed out that the major components of the feedstocks 
used herein are hydrocarbons with boiling points ranging from 25.degree. 
C. to 175.degree. C. including straight-chain and branched-chain paraffins 
and naphthenes, such as n-heptane, isooctane, and methyl cyclohexane, and 
aromatics such as BTX. 
The BTX fraction can include benzene, toluene, the C.sub.8 aromatics 
including ortho xylene, meta xylene, paraxylene, and ethyl benzene, and 
C.sub.9 aromatics, which, if present at all, appear in the smallest 
proportion in relation to the other components. 
The solvents used in solvent extraction/steam distillation processes for 
the recovery of BTX are water miscible organic liquids (at process 
temperatures) having a boiling point of at least about 200.degree. C. and 
having a decomposition temperature of at least about 225.degree. C. The 
term "water-miscible" includes those solvents which are completely 
miscible over a wide range of temperatures and those solvents which have a 
high partial miscibility at room temperature since the latter are usually 
completely miscible at process temperatures. The solvents are also polar 
and are generally comprised of carbon, hydrogen, and oxygen with some 
exceptions. Examples of solvents which may be used in the process of this 
invention are dipropylene glycol, tripropylene glycol, dibutylene glycol, 
tributylene glycol, ethylene glycol, diethylene glycol, ethylene glycol 
monomethyl ether, ethylene glycol monoethyl ether, diethylene glycol 
monomethyl ether, diethylene glycol monoethyl ether, sulfolane, N-methyl 
pyrrolidone, triethylene glycol, tetraethylene glycol, ethylene glycol 
diethyl ether, propylene glycol monoethyl ether, pentaethylene glycol, 
hexamethylene glycol, and mixtures thereof. The preferred group of 
solvents is the polyalkylene glycols and the preferred solvent is 
tetraethylene glycol. 
Additional solvents, which may be used alone or together, or with the 
aforementioned solvents are amides such as formamide, acetamide, 
dimethylformamide, diethylformamide, and dimethylacetamide; amines such as 
diethylenetriamine and triethylenetetramine; alkanolamines such as 
monoethanolamine, diethanolamine, and triethanolamine; nitriles such as 
beta,beta.sup.1 -oxydipropionitrile and beta,beta.sup.1 
-thiodipropionitrile; phenol and the cresols; the methyl sulfolanes; 
sulfoxides such as dimethyl sulfoxide and diethyl sulfoxide; lactones such 
as gamma propiolactone and gamma-butyrolactone. 
The apparatus used in the process both for extraction and distillation is 
conventional, e.g., an extraction column of the multi-stage reciprocating 
type containing a plurality of perforated plates centrally mounted on a 
vertical shaft driven by a motor in an oscillatory manner can be used as 
well as columns containing pumps with settling zones, sieve trays with 
upcomers, or even a hollow tube while the distillation can be conducted in 
a packed, bubble plate, or sieve tray fractionating column. 
Counter-current flows are utilized in both extraction and distillation 
columns. 
Heat exchangers, decanters, reservoirs, solvent regenerators, condensers, 
compressors, and pumps as well as various extractors other than the main 
extractor can also be used to complete the system. The other extractors 
are preferably single stage mixer-settlers, but can be any of the well 
known types. Again, all of this apparatus is conventional off the shelf 
equipment commonly used in extraction/distillation processes. 
The solvent is used as an aqueous solution containing water in an amount of 
about 1 percent to about 10 percent by weight based on the weight of the 
solvent and preferably containing water in an amount of about 2 percent to 
about 6 percent by weight. 
Generally, to accomplish the extraction, the ratio of solvent (exclusive of 
water) to feedstock in the extractor is in the range of about 4 to about 8 
parts by weight of solvent to one part by weight of feedstock. This broad 
range can be expanded upon where nonpreferred solvents are used. A broad 
range of about 3 to about 12 parts by weight of solvent to one part by 
weight of feedstock and a preferred range of about 5 parts to about 7 
parts of solvent per part of feedstock can be used successfully for the 
solvent of preference and other like solvents. In final analysis, however, 
the ratio is selected by the technician based on experience with the 
particular feedstock and depends in part upon whether high recovery or 
high purity is being emphasized. 
The reflux to the extraction zone, an important part of the process, is 
generally made up of about 20 percent to about 50 percent by weight 
aliphatics having from 5 to 7 carbon atoms and about 50 percent to about 
80 percent by weight aromatics, both based on the total weight of the 
reflux. The ratio of reflux to feedstock in the extraction zone is, 
generally, maintained in the range of about 0.5 to about 1.5 parts by 
weight of reflux to one part by weight of feedstock and preferably about 
0.5 to about 1.0 part by weight of reflux to one part by weight of 
feedstock, but, again, is selected by the technician just as the ratio of 
solvent to feedstock. The reflux aliphatics pass into the extract rather 
than being taken overhead with the raffinate and are recycled to the 
extractor from the reflux decanter. 
The temperature in the extraction zone is maintained in the range of about 
100.degree. C. to about 200.degree. C. and is preferably in the range of 
about 125.degree. C. to about 150.degree. C., especially for the solvent 
of preference. 
The pressure in the extraction zone is maintained in the range of about 75 
psig to about 200 psig. As is well know in the art, however, one selected 
pressure is not maintained throughout the extraction zone, but, rather, a 
high pressure within the stated range is present at the bottom of the zone 
and a low pressure, again within the stated range, is present at the top 
of the zone with an intermediate pressure in the middle of the zone. The 
pressures in the zone depend on the design of the equipment and the 
temperature, both of which are adjusted to maintain the pressure within 
the stated range. 
The temperature at the top of the distillation zone, which, in terms of the 
apparatus used, may be referred to as a distillation column or stripper, 
is at the boiling point of the mixture of aromatics present in the zone 
while the temperature at the bottom of the stripper is generally in the 
range of about 135.degree. C. to about 200.degree. C. 
The pressure at the top of the stripper, an upper flash zone in this case, 
is in the range of about 20 psig to about 45 psig. In a lower flash zone 
just beneath the upper flash zone and connected thereto, the pressure is 
in the range of about 10 psig to about 25 psig and is about 10 or 20 psig 
lower than the pressure in the upper flash zone. The pressure in the rest 
of the distillation zone is maintained in the range of about 15 psig to 
about 25 psig with some variation throughout the zone. 
The steam or steam/water mixture brought into the bottom of the 
distillation zone enters at a temperature of about 100.degree. C. to about 
150.degree. C. and is under a pressure of about 15 psig to about 25 psig. 
The total water and/or steam injected into the distillation column is in 
the range of about 0.1 part to about 0.5 part by weight of water to one 
part by weight of aromatics in the zone and preferably in the range of 
about 0.1 part to about 0.3 part by weight of water to one part by weight 
of aromatics. The water used for the stripping steam is usually called 
stripping water. A small amount of water is present in liquid form in the 
distillation zone dissolved in the solvent. 
Typically, in solvent extraction/steam distillation processes, the 
feedstock is preheated and then introduced to the main extractor at about 
the middle tray. An aqueous solvent solution (known as lean solvent) 
enters at the top tray of the extractor and percolates down the column 
removing aromatics from the feedstock. The raffinate, essentially free of 
aromatics, leaves the top of the column. Provisions are made for the 
recovery of solvent and any remaining aromatics from the raffinate as well 
as the water which is used to wash it. In the lower half of the extractor, 
the solvent solution of aromatics comes into countercurrent contact with a 
reflux liquid, which enters the extractor below the bottom tray. The 
reflux percolates up the lower half of the extractor progressively 
dissolving in and purifying the solvent solution of aromatics. The extract 
(known as rich solvent) leaves the bottom of the extractor and enters the 
stripper (or distillation zone) at an upper flash chamber. Part of the 
extract flashes on entering the flash chamber and is taken overhead in 
vapor form and the other part of the extract passes as a liquid into a 
lower flash chamber. Again, part of the extract, flashes overhead and the 
balance of the extract (at least about 80 percent by weight) percolates 
down the column into the fractionation zone where it comes into 
countercurrent contact with the stripping vapors, i.e., steam, and more 
vapors are generated. A part of the vapors rises to the top of the column 
where it mixes with flash vapors to form the overhead distillate. The 
overhead distillate provides reflux for the extractor. After the rich 
solvent descends about halfway down the column, it becomes essentially 
free of aliphatics. At this point, a vapor side-cut distillate is removed. 
The side cut distillate is separated into its aromatics and solvent/water 
components, the aromatics being recovered and the solvent and water being 
recycled into the system. Stripping water from the side cut distillate and 
other water from the system is returned to the bottom of the stripper as 
steam or a steam/water mixture. The bulk of the solvent and water leaves 
the bottom of the stripper. A portion of this solution is directed to a 
reboiler where it is vaporized and then returned to a point below the 
bottom tray of the stripper to provide heat therefor. The balance of the 
solvent/water solution is recycled to the top tray of the main extractor. 
There are many specific variations of the above process, each of which 
seeks either to reduce apparatus requirements, i.e., capital expenditure, 
or energy consumption, or make more effective use of process components 
while meeting purity specifications. 
SUMMARY OF THE INVENTION 
By this invention steam distillations can be conducted with reduced energy 
consumption by the indirect heat exchange of the overhead from the 
distillation zone with water at a pressure lower than that in the 
distillation zone. This lower pressure is sufficient to enable water to be 
vaporized and is maintained by passing the vaporized overhead to the low 
pressure port of a fluid ejector through which a higher pressure stream 
entering the distillation zone, such as a feed stream or steam, is passed. 
Additional reduction of energy consumption can be achieved by passing the 
unvaporized water from the heat exchanger to a motive steam generator and 
using the motive steam as the fluid for the fluid ejector. 
Not only can the processes of this invention offer reduced energy 
consumption, but, also, the processes enable the re-use of stripping 
water. In many steam distillation processes, the stripping water in the 
overhead from the distillation zone contains minor portions of the 
components to be separated. Thus, the stripping water may not be suitable 
for disposal or for use in other process equipment such as steam boilers. 
The processes of this invention can enable this stripping water to be 
recycled to the steam distillation zone in an economic and efficient 
manner in which heat is recovered from the overhead stream and effectively 
returned to the steam distillation zone. Moreover, the process may be 
practiced with little capital investment and without undue maintenance 
because of the use of the fluid ejector to maintain the lower pressure in 
the heat exchange. 
The process may be suitable for various steam distillation operations 
wherein a substantially water immiscible component is being separated. 
These separations include the separation of hydrocarbons, essential oils, 
fatty acids, turpentine, pine oil, camphor, monomers from polymers and the 
like, and can find application in processes such as acid gas removal 
processes, the Benfield process, alkanolamine acid gas treating system, 
and the like. 
According to one aspect of the invention, an improvement has been found in 
a steam distillation process for the recovery of hydrocarbons wherein 
there is (i) a primary flash zone at the top of the distillation zone in 
which rich solvent is flashed and/or (ii) provision for the removal of 
side-cut distillate vapors from about the middle of the distillation zone. 
The improvement comprises (a) heat exchanging flashed rich solvent vapors 
or side-cut distillate vapors with stripping water to provide stripping 
water vapors and stripping water at at least about the boiling point of 
water; (b) passing the stripping water vapors from step (a) to a steam 
ejector; (c) passing the stripping water from step (a) to a motive steam 
generator wherein the stripping water is vaporized under pressure; (d) 
passing the stripping water vapors from step (c) to the steam ejector 
referred to in step (b); and (e) passing the stripping water vapors, 
introduced into the steam ejector in accordance with steps (b) and (d), to 
the lower half of the distillation zone. 
In another aspect of the invention, an overhead vapor stream from the 
distillation zone is heat exchanged with stripping water at a temperature 
which under the pressure of the heat exchanging is at least about the 
boiling point of water whereby stripping water vapors are produced and 
passed to a steam ejector. Steam, at a higher pressure, is passed through 
the steam ejector into the distillation zone whereby the pressure of the 
heat exchange is lower than the pressure of the steam distillation. 
In a further aspect of the invention, a feed stream containing at least one 
substantially water immiscible component to be separated and an operative 
stream containing at least one of water and steam are introduced into a 
steam distillation vessel which is maintained under steam distillation 
conditions including temperature and pressure to provide a vaporous 
overhead stream containing the at least one component to be separated and 
a liquid bottoms fraction. The liquid bottoms fraction is withdrawn from a 
lower portion of the vessel and the overhead stream is withdrawn from an 
upper portion of the vessel and is passed through an indirect heat 
exchanger. The overhead stream is condensed to provide a liquid stream 
rich in the at least one component to be separated and a water stream. At 
least a portion of the water stream is passed to the indirect heat 
exchange as the heat exchange medium and at least a portion of the water 
stream in the indirect heat exchanger is vaporized at a lower absolute 
pressure than the pressure in the steam distillation vessel. This 
vaporized stream is passed to a lower pressure inlet of a fluid ejector 
through which at least one of the feed stream and at least a portion of 
the operative stream is passed at a higher absolute pressure into the 
steam distillation vessel whereby the heat exchange medium side of the 
indirect heat exchanger is maintained at said lower absolute pressure 
sufficient to generate steam using heat contained in the overhead stream.

DETAILED DESCRIPTION 
The main extractor, feedstock, solvent, temperatures, and pressures are as 
described above except as noted. While subject process can be applied to 
any steam distillation process, which provides for a primary flash zone 
and/or side cut distillate vapors, the application of particular interest 
is a solvent extraction/steam distillation process for the recovery of 
aromatic hydrocarbons. 
Referring to the drawing: 
The rich solvent from the extractor (not shown) is at a temperature in the 
range of about 100.degree. C. to about 150.degree. C. It passes along line 
1 to primary flash chamber 2 at the top of stripper 3. Primary flash 
chamber 2 is maintained at a pressure in the range of about 20 pounds per 
square inch gauge (psig) to about 60 psig. Part of the hydrocarbon and 
water in the rich solvent is flashed overhead along line 4 to pass as a 
vapor through line 5 at a temperature in the range of about 90 to about 
140.degree. C. and at a pressure in the range of about 15 psig to about 55 
psig entering, prior to its condensation, into stripping water vaporizer 
(heat exchangers) 6. The stripping water enters vaporizer 6 along line 7 
at a temperature in the range of about 35.degree. to about 80.degree. C. 
As the primary flash vapors condense, the stripping water is heated to 
about 100.degree. C., the boiling point of water at atmospheric pressure. 
Part of the stripping water is vaporized at about one atmosphere. The 
other part remains as a liquid. In the event that vaporizer 6 is operated 
at less than atmospheric pressure, the boiling point of water will, of 
course, be reduced accordingly. 
The stripping water is split into two streams, the vapor following line 9 
and the liquid following line 10. The condensed primary flash vapors 
proceed along line 5 where they meet vapors from secondary flash chamber 
11 and the top of stripper 3 passing along lines 12 and 13, respectively, 
and combining into line 14. Streams 5, 12, and 13 represent the overhead 
distillate. Streams 5 and 14 combine and enter stream 15, which is 
introduced into reflux condenser 16. The vapors are condensed in reflux 
condenser 16 and the liquid passes into decanter 17 where a hydrocarbon 
reflux phase is separated from a water phase. The reflux is recycled to 
the extractor and the water phase is combined with the water phase from 
decanter 29 and sent to pump 18 for reuse as stripping water. The water 
phase from decanter 17 is passed along line 17A and the water phase from 
decanter 29 is passed along line 29A. The stripping water, which is at a 
temperature in the range of about 35.degree. to about 80.degree. C. is 
passed from pump 18 along line 7 to stripping water vaporizer 6 as noted 
above. 
The stripping water, at about 100.degree. C., passes through line 10 to 
pump 19 and thence to motive steam generator 20 where it is converted to 
high pressure steam with a temperature in the range of about 170.degree. 
to about 230.degree. C. and at a pressure in the range of about 100 to 
about 400 psig This is accomplished by introducing steam at a pressure in 
the range of about 125 to about 450 psig along line 21 into motive steam 
generator 20. The stripping water steam (or motive steam) from generator 
20 then passes along line 22 to steam ejector 23 providing the driving 
force therefor. The stripping water vapor at 100.degree. C. enters steam 
ejector 23 along line 9 and is pumped into stripper 3. Essentially all of 
the steam from steam ejector 23 is pumped into stripper 3. 
The content of solvent in the stripping water entering generator 20 is less 
than about one percent by weight. This small amount of solvent 
concentrates in generator 20 and is purged out of generator 20 and into 
stripper 3 by using a purge stream not shown in the drawing. 
The steam used in generator 20 continues along line 21 into reboiler 24 
where it vaporizes a portion of the lean solvent/water solution passing 
along line 25 from the bottom of stripper 3. The steam is condensed and 
leaves the system along line 21 while the lean solvent/water solution 
vapor is returned to stripper 3 along line 25. The bulk of the lean 
solvent/water solution from the bottom of stripper 3 passes along line 26 
to the top of the main extractor. 
The side-cut distillate vapors pass from the middle of stripper 3 through 
line 27 to condenser 28. The now liquid side-cut distillate then passes 
into decanter 29 where an aromatics phase is separated from a water phase. 
The water phase is recycled as stripping water to pump 18 and the 
aromatics phase is recovered for further distillation and separation. 
An alternate procedure (not shown) is to use the side-cut distillate vapors 
instead of the primary flash vapors. The side-cut distillate vapors, at a 
temperature in the range of about 90.degree. to about 140.degree. C. and a 
pressure of about 0 psig to about 20 psig, are introduced into stripping 
water vaporizer 6. The procedure, then is the same as described for the 
primary flash vapors. 
After the heat is obtained from the side-cut distillate vapors, the 
remaining vapors pass to condenser 28 and the condensate then continues 
along line 27. Further, the alternate procedures can be combined, i.e., 
the heat can be recovered form both the primary flash vapors and the 
side-cut distillate vapors. To accomplish this, an additional stripping 
water vaporizer is needed for the side-cut distillate vapors together with 
additional piping to complete the scheme. The key to the energy recovery 
is using the primary flash vapors and/or side cut distillate vapors before 
the vapors expand, i.e., while they are under pressure, the pressure being 
in the range of about 20 to about 60 psig for the primary flash vapors and 
about 0 to about 25 psig for the side cut distillate vapors. In order of 
preference, i.e., achieving the highest heat recovery, the side-cut 
distillate vapors appears to be first, the use of both primary flash 
vapors and side cut distillate vapors, second, and the primary flash 
vapors, third. This order can change, however, depending on the particular 
case to which the invention is applied. The recovery of heat from the 
vapors is enhanced by the use of a high flux tubing heat exchanger, which 
make temperature approaches of about 2.degree. to about 3.degree. C. 
feasible. The purity of the side-stream distillate vapors makes the 
stripping water vaporizer a good candidate for a high flux tubing 
application. 
The advantages of subject process are as follows: 
1. High energy savings. Further, the higher the stripping water rate used 
to strip the aromatics, i.e., the higher the aromatic content of the feed, 
the greater the energy savings obtained. 
2. The process is applicable to any distillation column that uses stripping 
water to remove hydrocarbons (or any other solute) from a solvent. 
3. The cost of the stripping water vaporizer and the motive stream 
generator are offset by the elimination of other heat exchangers required 
in comparable systems. 
4. Steam ejectors are inexpensive as compared to the usual compressors. 
The invention is illustrated by the following example (percentages and 
ratios are by weight): 
The process described above and in the drawing is carried out twice in the 
preferred mode, once using the primary flash vapors (process A) and the 
other time using the side-cut distillate vapors (process B). The feedstock 
is characterized as a high severity reformate containing about 63 percent 
BTX. The lean solvent solution contains about 94 percent tetraethylene 
glycol and about 6 percent water. 
The operating conditions and results are the same for process A and process 
B except as noted. They are as follow: 
temperature of rich solvent entering stripper 3: 138.degree. C. 
pressure in primary flash chamber: 35 psig 
temperature of primary flash vapors: 129.degree. C. 
temperature of side-cut distillate vapors: 126.degree. C. 
pressure of side cut distillate vapors (before expansion): 10 psig 
temperature in stripper 3: 156.degree. C. 
pressure in stripper 3: 12 psig 
temperature of stripping water vapors in line 9: 100.degree. C. 
pressure of stripping water vapors in line 9: 1 atmosphere 
temperature of stripping water in line 7: 49.degree. C. 
temperature of stripping water in line 10: 100.degree. C. 
pressure of steam entering line 21: 200 psig 
pressure of motive steam in line line 22: 125 psig 
feedstock rate (pounds per hour): 116,198 
solvent solution to feedstock ratio: 5.2 
reflux to feedstock ratio: 0.78 
stripping water rate (pounds per hour): 29,336 
primary flash vapors (pounds per hour): 17,341 
side-cut distillate vapors (pounds per hour): 92,747 
Recoveries, i.e., percent of recovery based on amount in feedstock: 
benzene: 99.97 
toluene: 99.78 
xylene: 98.55 
cumene: 84.48 
Impurities (parts per million by weight): 632 
reboiler duty for Process A (10.sup.6 BTU's per hour) 55.1 
reboiler duty for Process B (10.sup.6 BTU's per hour) 51.0 
estimated energy saved in Process A (10.sup.6 BTU's per hour). 8.15 
estimated energy saved in Process B (10.sup.6 BTU's per hour) 12.0 
estimated energy reduction in Process A (percent): 12 
estimated energy reduction in Process B (percent): 19 
FNT Note: Energy savings and percentage reduction are based on a comparison 
with a process run using the same steps and conditions except that the 
primary flash vapors and side-cut distillate are not used to heat the 
stripping water. Instead a rich solvent/stripping water heat exchanger is 
used to provide heat for the stripping water.