Multistage process for converting olefins or oxygenates to heavier hydrocarbons

A multistage technique for converting olefins to heavier hydrocarbons including a sorption prefractionation unit for separating olefinic feedstock into a sorbate stream rich in liquified olefins and a vapor stream rich in light olefins; a first stage catalytic reactor unit for oligomerizing olefins from the sorbate stream including means for maintaining the first stage at elevated pressure and predetermined temperature for producing substantially linear aliphatic hydrocarbons; a second stage catalytic reactor unit for oligomerizing light olefin including means for maintaining the second stage under high severity conditions at substantially higher temperature than the first stage; and a product fractionation unit for separating effluent from the first and second stages to separate and recover heavy hydrocarbon product and a sorbent recycle fraction. The sorbent fraction is recycled to the sorption prefractionation unit for contacting olefinic feedstock with the recycled sorbent. Unconverted second stage olefins may be also sorbed by the recycled fraction for further reaction.

FIELD OF THE INVENTION 
This invention relates to a system for converting light olefins or organic 
oxygenate precursors, such as methanol or dimethyl ether (DME), to liquid 
hydrocarbons. In particular it provides a continuous process for producing 
distillate range fuel products by oligomerizing fractionated olefins to 
produce a major amount of distillate product for use as diesel fuel or the 
like. 
BACKGROUND OF THE INVENTION 
Recent developments in zeolite catalysts and hydrocarbon conversion 
processes have created interest in utilizing olefinic feedstocks, for 
producing C.sub.5.sup.+ gasoline, diesel fuel, etc. In addition to the 
basic work derived from ZSM-5 type zeolite catalysts, a number of 
discoveries have contributed to the development of a new industrial 
process, known as Mobil Olefins to Gasoline/Distillate ("MOGD"). This 
process has significance as a safe, environmentally acceptable technique 
for utilizing feedstocks that contain lower olefins, especially C.sub.2 
-C.sub.5 alkenes. This process may supplant conventional alkylation units. 
In U.S. Pat. Nos. 3,960,978 and 4,021,502, Plank, Rosinski and Givens 
disclose conversion of C.sub.2 -C.sub.5 olefins, alone or in admixture 
with paraffinic components, into higher hydrocarbons over crystalline 
zeolites having controlled acidity. Garwood et al have also contributed 
improved processing techniques to the MOGD system, as in U.S. Pat. Nos. 
4,150,062, 4,211,640 and 4,227,992. The above-identified disclosures are 
incorporated herein by reference. 
Conversion of lower olefins, especially propene and butenes, over HZSM-5 is 
effective at moderately elevated temperatures and pressures. The 
conversion products are sought as liquid fuels, especially the 
C.sub.5.sup.+ aliphatic and aromatic hydrocarbons. Olefinic gasoline can 
be produced in good yield by the MOGD process and may be recovered as a 
product or recycled to the reactor system for further conversion to 
distillate-range products. Distillate mode operation can be employed to 
maximize production of C.sub.10.sup.30 aliphatics by reacting the lower 
olefins at high pressure and moderate temperature. Operating details for 
typical MOGD units are disclosed in copending U.S. Pat. applications Ser. 
No. 488,834, filed Apr. 26, 2983 (Owen et al), now U.S. Pat. No. 4,456,779 
and Ser. No. 481,704, filed Apr, 4, 1983 (Tabak), incorporated herein by 
reference. 
In addition to their use as shape selective oligomerization catalysts, the 
medium pore ZSM-5 type catalysts are useful for converting methanol and 
other lower aliphatic alcohols and/or corresponding ethers to olefins. 
Particular interest has been directed to a catalytic process for 
converting low cost methanol to valuable hydrocarbons rich in ethene and 
C.sub.3.sup.+ alkenes. Various processes are described in U.S. Pat. Nos. 
3,894,107 (Butter et al), 3,928,483 (Chang et al), 4,025,571 (Lago), and 
in copending U.S. patent application Ser. No. 388,768, filed June 15, 1982 
(Yurchak et al), now abandoned. Significance of the methanol-to-olefins 
("MTO") type processes, especially for producing ethene, is discussed in 
Hydrocarbon Processing, Nov. 1982, pp. 117-120. It is generally known that 
the MTO process can be optimized to produce a major fraction of C.sub.2 
-C.sub.4 olefins. 
SUMMARY OF THE INVENTION 
A continuous catalytic process has been found for converting olefinic 
feedstock comprising ethylene and C.sub.3.sup.+ olefins to heavier liquid 
hydrocarbon product. Methods and means are provided for prefractionating 
the olefinic feedstock to obtain a gaseous stream rich in ethylene and a 
liquid stream containing C.sub.3.sup.+ olefin; heating and contacting the 
liquid stream from the prefractionating step with shape selective medium 
pore zeolite oligomerization catalyst in a distillate mode catalytic 
reactor system at elevated temperature and pressure to provide a heavier 
hydrocarbon effluent stream comprising distillate, gasoline and lighter 
hydrocarbons; fractionating the effluent stream to recover distillate, 
gasoline and lighter hydrocarbons; recycling at least a portion of the 
recovered gasoline as a first liquid sorbent stream to the 
prefractionating step; further reacting the recycled gasoline together 
with sorbed C.sub.3.sup.+ olefin in the distillate mode catalytic reactor 
system; combining the ethylene-rich gaseous stream with a liquid 
hydrocarbon stream containing heavier liquid hydrocarbons; contacting the 
combined ethylene-rich stream at elevated temperature and pressure in a 
high severity reaction zone with shape selective medium pore zeolite 
oligomerization catalyst to convert at least a portion of the olefinic 
components to heavier hydrocarbons; cooling oligomerization reaction 
effluent from the high severity reaction zone to condense at least a 
portion of the heavier hydrocarbons; separating the cooled and partially 
condensed high severity reactor effluent stream into a vapor stream 
comprising unreacted light olefin and a condensed liquid hydrocarbon 
stream; fractionating the condensed liquid hydrocarbons from the high 
severity reaction effluent to provide a recycle sorbent stream and at 
least one product hydrocarbon stream; contacting the vapor stream from the 
high severity reaction effluent under sorption pressure conditions with 
cooled recycle sorbent stream to sorb said unreacted light olefin into the 
sorbent stream; and recycling the sorbent stream rich in olefin for 
further conversion with the olefinic feedstock. Advantageously, the first 
and second sorbent streams are recovered in a common fractionation unit. 
The oligomerization catalyst preferably comprises ZSM-5 type zeolite. Other 
objects and features of the invention will be seen in the following 
description and drawings.

DESCRIPTION OF PREFERRED EMBODIMENTS 
Catalyst versatility permits the same zeolite to be used in both the 
distillate mode primary stage and high severity secondary oligomerization 
stage. While it is within the inventive concept to employ substantially 
different catalysts in these stages, it is advantageous to employ a 
standard ZSM-5 having a silica alumina molar ratio of 70:1. 
The oligomerization catalysts preferred for use herein include the medium 
pore shape selective crystalline aluminosilicate zeolites having a silica 
to alumina ratio of at least 12, a constraint index of about 1 to 12 and 
acid cracking activity of about 160-200. Representative of the ZSM-5 type 
zeolites are ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35 and ZSM-38. 
ZSM-5 is disclosed in U.S. Pat. No. 3,702,886 and U.S. Pat. No. Re. 
29,948. Other suitable zeolites are disclosed in U.S. Pat. Nos. 3,709,979; 
3,832,449, 4,076,979, 3,832,449, 4,076,842, 4,016,245 and 4,046,839, 
4,414,423 and 4,417,086-8. The disclosures of these patents are 
incorporated herein by reference. A suitable catalyst for each fixed bed 
operation consisting essentially of HZSM-5 zeolite with 35 wt. % alumina 
binder in the form of cyclindrical extrudates of about 1-5 mm diameter. 
Other catalysts which may be employed for converting methanol/DME to lower 
olefins and oligomerizing same include the borosilicate, ferrosilicate, 
"silicalite" and/or synthetic mordenite materials. 
In this description, metric units and parts by weight are employed unless 
otherwise stated. While various reactor configurations may be used, 
including fluidized bed catalytic reactors, moving bed and fixed bed 
reactors, the invention is described for use in a plurality of fixed bed 
reactors operated under differing process conditions depending upon 
relative position in the system. 
Referring to FIG. 1, olefinic feedstock is contacted with a gasoline 
sorbent stream in a prefractionation section to provide a C.sub.3.sup.+ 
liquid stream, which is passed to the distillate mode oligomerization 
reactor system. The combined olefinic stream (containing recycled olefinic 
gasoline rich in C.sub.3 -C.sub.4 olefin sorbate) is reacted at high 
pressure and elevated temperature over the oligomerization catalyst. 
Distillate-rich effluent is then passed to a product fractionation system 
and separated into light gases (LPG), C.sub.5.sup.+ gasoline for recycle 
and distillate range hydrocarbons. The distillate stream contains a major 
amount of high boiling aliphatics, such s C.sub.10 -C.sub.20 hydrocarbons, 
and a minor amount of aromatics. The distillate product may be further 
stabilized by hydrotreating (HDT) in a relatively mild process to saturate 
the olefinic compounds and convert aromatics to corresponding naphthenes 
without substantial cracking or dealkylation to yield a distillate fuel 
product. Ethylene (ethene, C.sub.2 H.sub.2) recovered from the 
prefractionator is converted in a high severity reactor system, the 
effluent from which is cooled to condense C.sub.5.sup.+ hydrocarbons for 
further fractionation. For this purpose the condensed effluent from the 
high severity reactor may be combined with the distillate-rich stream. 
Unconverted vapor effluent may be recovered by absorption in a liquid 
sorbent, such as recycled gasoline, and further reacted in the high 
severity reactor. Since the recycled streams are liquid, they may be 
repressurized at moderate cost by pumping. 
In the process for catalytic conversion of olefins to heavier hydrocarbons 
by catalytic oligomerization using an acid crystalline zeolite, such as 
ZSM-5 type catalyst, process conditions can be varied to favor the 
formation of either gasoline or distillate range products. At moderate 
temperature and relatively high pressure, the conversion conditions favor 
distillate range product having a normal boiling point of at least 
165.degree. C. (330.degree. F.). Lower olefinic feedstocks containing 
C.sub.2 -C.sub.6 alkenes may be converted selectively; however, the 
distillate mode conditions do not convert a major fraction of ethylene. 
While propene, butene-1 and others may be converted to the extent of 50 to 
95% in the distillate mode, only about 10 to 50% of the ethylene component 
will be consumed. Accordingly, the ethene is advantageously recovered 
prior to the oligomerization stage. In a preferred embodiment the olefinic 
feedstock is obtained from lower aliphatic oxygenates, such as methanol or 
dimethylether (DME). FIG. 2 depicts a fixed bed multi-reactor (MTO) system 
for converting methanol (CH.sub.3 OH) and/or DME. A typical crude methanol 
feedstock may contain 4 to 17% water, with minor amounts of carbon oxides, 
methane, DME, etc. 
In the oxygenate conversion stage, ethene production may be optimized by 
employing fixed bed primary stage conditions in the temperature range of 
about 260.degree. C. to 425.degree. C., a pressure range of about 170 to 
800 kPa and weight hourly space velocity range of about 0.5 to 1.0 based 
on ZSM-5 equivalent catalyst and methanol equivalent in the primary stage 
feedstock. Typically about 25 to 90% of oxygenate feed is converted per 
reactor pass and water diluent is cofed with methanol and/or dimethyl 
ether in a molar ratio of about 0.1:1 to 5:1. Under these conditions, the 
primary stage hydrocarbon effluent usually contains about 25 to 40 wt. % 
ethene, about 10 to 50 wt. % propene, about 2 to 30 wt. % butene, less 
than 10 wt. % C.sub.1 to C.sub.4 paraffins, and about 5 to 20 wt. % 
aromatics, including about 1 to 5 wt. % durene. 
In the embodiment of FIG. 2, the feedstock is methanol (MeOH), which may be 
partially dehydrated in a separate process step over gamma-alumina 
catalyst to yield dimethyl ether (DME) and water. A preliminary dewatering 
step can be used to provide a feedstock consisting essentially of MeOH 
and/or DME; however, the presence of water in the MTO reactor may be 
beneficial. The oxygenate is fed continuously under low pressure through 
line 1 and heat exchanger 2, where it is raised to process temperature, 
and introduced to the first stage MTO reactor system. The initial 
dehydration reactor 10 is followed by a series of fixed bed catalytic 
reactors 11, 12, 13, 14 and 15 containing zeolite conversion catalyst. 
Inter-reactor coolers 10A, 11A, 12A, 13A, 14A and effluent cooler 15A 
control temperature of the process stream. Interstage separation means is 
provided for recovering water and hydrocarbons from the primary stage 
effluent stream. The first stage effluent is cooled to condense water and 
a major amount of C.sub.5.sup.+ liquid hydrocarbons. These liquids are 
separated from the hydrocarbon vapor in phase separator means 16. 
Byproduct water may be recovered from unreacted feedstock and discarded, 
or optionally a portion may be recycled, as indicated by the dashed lines. 
The liquid hydrocarbon phase and the ethene-rich light hydrocarbon vapor 
streams are recovered from separator 16. Optionally, unconverted DME may 
be removed by absorber 20. This optional section includes a compressor 19 
and stripper 21. Other suitable fixed bed catalytic processes for 
conversion of methanol/DME to lower olefins are described in U.S. Pat. 
Nos. 4,387,263, 4,393,265, 4,361,715 and South African Patent Application 
Ref. No. V01750 (Clover et al) filed Mar. 30, 1983, the entire disclosures 
of which are incorporated herein by reference. While the primary stage 
dehydration reactor has been exemplified herein by a fixed bed unit, a 
suitable fluid catalyst apparatus is disclosed in U.S. Pat. No. 4,379,123 
(Daviduk and Haddad). 
A typical MTO operation is conducted over a fixed bed of small crystal 
(0.02-0.05.mu.) HZSM-5/alumina extrudate catalyst at about 170 kPa (25 
psia), with a 1:1 H.sub.2 O:CH.sub.3 OHequivalent ratio at 315.degree. C. 
(600.degree. F.) at a space velocity (WHSV=0.5-1) to convert about 50% of 
the oxygenated organic feedstock components to hydrocarbons. Table I lists 
the organic hydrocarbon product distribution from a typical MTO process. 
TABLE I 
______________________________________ 
MTO Product Distribution 
Component wt. % 
______________________________________ 
Methane, wt. % 0.6 
Ethylene, wt. % 26.2 
Ethane, wt. % 0.1 
Propylene, wt. % 22.8 
Propane 3.9 
Butenes 7.9 
Isobutane 3.9 
n-Butane 2.6 
Pentenes 2.4 
C.sub.5 P + N 7.1 
C.sub.6 P + N 5.1 
C.sub.7 O 0.6 
C.sub.7 P + N 3.2 
C.sub.7 O 0.7 
C.sub.8 P + O + N 
2.1 
C.sub.9 P + O + N 
1.3 
C.sub.10 P + O + N 
1.1 
Benzene 0.1 
Toluene 0.5 
C.sub.8 Aromatics 
3.5 
C.sub.9 Aromatics 
2.1 
C.sub.10 Aromatics 
2.2 
(Durene) (1.7) 
______________________________________ 
SORPTION FRACTIONATION 
The light hydrocarbon stream recovered from the primary conversion stage 
preferably contains a major amount of C.sub.2 -C.sub.4 olefins. The novel 
system includes a first fractionation means for recovering ethene from the 
primary stage olefinic vapor including a sorption tower operatively 
connected to selectively sorb C.sub.3.sup.+ hydrocarbons from the olefinic 
vapor in a liquid sorption stream. Since the interstage fractionation unit 
is usually operated at a pressure higher than the primary stage and lower 
than the secondary conversion stage, vapor compression means for the 
primary stage light hydrocarbon stream and means for pressurizing and 
heating the liquid sorption stream containing C.sub.3.sup.+ sorbate are 
provided. 
A suitable sorption fractionation system is described in copending U.S. 
patent application Ser. No. 508,779 filed June 29, 1983 (Hsia et al), the 
disclosure of which is incorporated herein by reference. The C.sub.2.sup.- 
and C.sub.3.sup.+ separation is accomplished by a single absorber-stripper 
using gasoline recycle as absorbent and pumparounds for removing 
absorption heat. The amount of absorbent is set by the amount of recycle 
gasoline required in the C.sub.3.sup.+ olefins conversion reaction thereby 
allowing the tower bottom stream to be pumped directly to the reactor 
pressure. Without using refrigeration, this tower efficiently and 
effectively separates the ethylene and light gases (H.sub.2, CO, CO.sub.2 
and CH.sub.4) from the C.sub.3.sup.+ hydrocarbon. 
The gasoline sorbent is an aliphatic hydrocarbon mixture boiling in the 
normal gasoline range of about 50.degree. to 165.degree. C. (125.degree. 
to 330.degree. F.), with minor amounts of C.sub.4 -C.sub.5 alkanes and 
alkenes. Preferably, the total gasoline sorbent stream to feedstock weight 
ratio is greater than about 3:1; however, the content of C.sub.3.sup.+ 
olefinic components in the feedstock is a more preferred measure of 
sorbate to sorbent ratio. Accordingly, the process may be operated with a 
mole ratio of about 0.2 moles to about 10 moles of gasoline per mole of 
C.sub.3.sup.+ hydrocarbons in the feedstock, with optimum operation 
utilizing a sorbent:sorbate molar ratio about 1:1 to 1.5:1. 
It is understood that the various process conditions are given for a 
continuous system operating at steady state, and that substantial 
variations in the process are possible within the inventive concept. In 
the detailed examples, metric units and parts by weight are employed 
unless otherwise specified. 
Referring to FIG. 3 of the drawing, olefinic feedstock is introduced to the 
system through a feedstock inlet 101, passing through knockout pot 102 to 
remove any condensate, such as water. At this point the vapor is at a 
temperature from ambient up to about 38.degree. C. (100.degree. F.) and a 
pressure of about 60 kPa (75 psig), and is pressurized by compressor means 
103 to about 2230 kPa (310 psig). This pressurized stream is then mixed 
with a liquid downcomer stream 104 from fractionating tower 120, cooled in 
exchanger 105 to about 35.degree.-40.degree. C. and passed to phase 
separator 106 where condensed water is removed. The hydrocarbon stream is 
then fed via conduit 107 connected between stages of fractionating 
sorption tower 120 wherein gaseous olefinic feedstock is contacted with 
liquid sorbent in a vertical fractionation column operating at least in 
the upper portion thereof in countercurrent flow. Effectively this unit is 
a C.sub.2 /C.sub.3.sup.+ splitter. The sorption tower employs a plate 
column; however, the fractionation equipment may employ vapor-liquid 
contact means of various designs in each stage including packed beds of 
Raschig rings, saddles or other porous solids or low pressure drop contact 
devices. 
Sorption tower 120, as depicted, has multiple contact zones, with the heat 
of adsorption being removed via interstage pump around cooling means 121A, 
B, and C. The liquid gasoline sorbent is introduced to the sorption tower 
through an upper inlet means 125 above the top contact section 120T. It is 
preferred to mix incoming liquid sorbent with outgoing splitter overhead 
ethylene-rich gas. High purity ethylene is recovered from the system 
through gas outlet 131 and sent to storage, further processing or 
conversion to other products. Liquid sorbent from separator 130 is then 
pumped to the upper liquid inlet 124 for countercurrent contact in a plate 
column or the like with upwardly flowing ethylene rich vapors. Liquid from 
the bottom of upper contact zone is pumped to a heat exchanger in loop 
121A, cooled and returned to the tower, then cooled in loop 121B adjacent 
to an intermediate contact zone, again cooled in loop 121C, and returned 
to the tower above contact zone 126, which is located above the feedstock 
inlet 107. Under tower design conditions of about 2060 kPa (300 psia), it 
is preferred to maintain liquid temperature of streams entering the tower 
from 121, 107 and 124 at about 40.degree. C. (100.degree. F.). The lower 
contact zone provides further fractionation of the olefin-rich liquid. 
Heat is supplied to the sorption tower by removing liquid from the bottom 
via reboiler loop 128, heating this stream in a heat exchanger, and 
returning the reboiled bottom stream to the tower below contact zone 126. 
The liquid sorbate-sorbent mixture is withdrawn through bottom outlet 129 
and pumped to storage or directly to the secondary stage for further 
reaction. The fractionator bottoms stream 129 is recovered at about 
120.degree. C. (250.degree. F.), then pumped to the higher reactor 
pressure (e.g., about 4670 kPa, 665 psig) and passed to the secondary 
conversion stage. A typical sorption fractionation material balance is 
given below, for steady state operation using olefinic feed gas from the 
MTO primary stage. The units are expressed in moles per hour. 
TABLE II 
______________________________________ 
SORPTION FRACTIONATION MATERIAL BALANCE 
GAS LIQUID 
FEED RECYCLE PRODUCT SORBATE 
GAS GASOLINE STREAM STREAM 
______________________________________ 
H.sub.2 
1.17 -- 1.17 -- 
CO 0.22 -- 0.22 -- 
CO.sub.2 
0.38 -- 0.38 -- 
C.sub.1 
0.63 -- 0.63 -- 
C.sub.2.spsb.= 
14.83 -- 14.78 0.05 
C.sub.2 
0.11 -- 0.10 0.01 
C.sub.3.spsb.= 
8.09 -- 0.36 7.73 
C.sub.3 
1.62 -- 0.02 1.60 
iC.sub.4 
1.00 0.09 0.03 1.06 
C.sub.4.spsb.= 
2.40 0.03 0.01 2.42 
nC.sub.4 
0.32 0.17 0.04 0.45 
iC.sub.5 
0.67 2.64 0.25 3.06 
C.sub.5 
0.24 1.23 0.11 1.36 
nC.sub.5 
0.06 0.22 0.02 0.26 
C.sub.6.sup.+ 
0.81 11.19 0.13 11.87 
H.sub.2 O 
0.37 -- 0.03 -- 
32.92 15.58 18.27 29.88 
______________________________________ 
DISTILLATE MODE 
Oligomerization Reactor Operation 
The main distillate production stage provides catalytic oligomerization 
reactor means containing medium pore shape selective zeolite 
oligomerization catalyst for converting olefinic hydrocarbons in the 
sorption stream to liquid hydrocarbons comprising a major amount of 
distillate. This process stream is passed to second fractionation means 
for separating secondary stage effluent into a light hydrocarbon stream 
rich in LPG (C.sub.3 -C.sub.4) aliphatic hydrocarbons, a C.sub.5.sup.+ 
gasoline stream and distillate range stream. Means is provided for 
recycling at least a portion of the C.sub.5.sup.+ gasoline stream to the 
first sorption fractionation means as a lean sorbent stream. 
A typical distillate mode secondary stage reactor system 220 is shown in 
FIG. 4. A plural reactor system may be employed with inter-reactor 
cooling, whereby the reaction exotherm can be carefully controlled to 
prevent excessive temperature above the normal moderate range of about 
190.degree. to 315.degree. (375.degree.-600.degree. F.). The olefinic 
feedstream comprising the C.sub.3.sup.+ light hydrocarbons and sorbent 
gasoline is introduced through conduit 229 and carried by a series of 
conduits through heat exchangers 212A, B, C and furnace 214 where the 
feedstream is heated to reaction temperature. The olefinic feedstream is 
then carried sequentially through a series of zeolite beds 220A, B, C 
wherein at least a portion of the olefin content is converted to heavier 
distillate constituents. Advantageously, the maximum temperature 
differential across only one reactor is about 30.degree. C. 
(.DELTA.T.about.50.degree. F.) and the space velocity (LHSV based on 
olefin feed) is about 0.5 to 1.5. The heat exchangers 212A and 212B 
provide inter-reactor cooling and 212C reduces the effluent to separation 
temperature. An optional heat exchanger 212D may further recover heat from 
the effluent stream 121 prior to separation. 
Preferably, the secondary stage process conditions are optimized to produce 
heavy liquid hydrocarbons having a normal boiling above 165.degree. C. 
(330.degree. F.) are fed as a continuous stream to a final fractionator 
unit (such as a distillation system). Gasoline, rich in C.sub.5.sup.+ 
olefins and lighter hydrocarbons are fractionated in a tower to provide an 
olefinic gasoline stream for recycle to the MOGD reactor system or 
recovered as product. The lighter hydrocarbons, rich in C.sub.3 -C.sub.4 
alkenes may be condensed and recovered as LPG product or optionally 
recycled to the MOGD reactor system. The secondary stage typical HZSM-5 
fixed bed reactor system depicted in FIG. 4, operates at 0.6 liquid hourly 
space velocity (based on olefins fed to reactors), 1:1 gasoline:olefin 
recycle ratio, temperature of 230.degree. C. (450.degree. F.) (SOC) to 
315.degree. C. (600.degree. F.) (EOC) and a total pressure of 4225 kPa 
(600 psig) at minimum olefin partial pressure at the inlet of 1100 kPa 
(160 psig). The secondary stage effluent from such a typical system is 
shown below. 
TABLE III 
______________________________________ 
Component Wt. % 
______________________________________ 
CH.sub.4 0.07 
C.sub.2 H.sub.6 0.13 
C.sub.3 H.sub.8 2.80 
IC.sub.4 H.sub.10 2.00 
NC.sub.4 H.sub.10 2.00 
i-C.sub.5 H.sub.12 
0.42 
n-C.sub.5 H.sub.12 
0.03 
C.sub.5 H.sub.10 0.95 
C.sub.6 -330.degree. (Gasoline) 
12.60 
330.degree. + (Distillate) 
79.00 
______________________________________ 
A typical product fractionation system is described in copending U.S. 
patent application Ser. No. 488,834 filed Apr. 26, 1983 (Owen et al), 
incorporated herein by reference. 
It is within the inventive concept to cascade substantially all 
C.sub.3.sup.+ vapor and liquid hydrocarbon product from the MTO stage into 
the distillate mode reactor followed by hydrotreating of the distillate 
product as depicted in FIG. 1. This will minimize the number of process 
steps and will maximize distillate production by polymerizing gasoline 
range olefins, and by alkylating gasoline range aromatics. Durene can be 
reduced via saturation to its corresponding naphthene in a subsequent mild 
hydrotreating step. Substantially all of the polymethylbenzenes or other 
aromatics formed in the dehydration reactor stage can be accumulated in 
the distillate fraction according to the present invention, and 
hydrotreated distillate durene content is decreased substantially below 2 
wt. %, preferrably below 1%. 
The present process is particularly useful in producing a major product 
stream wherein the 165 C.sup.+ fraction consists mainly of C.sub.10 to 
C.sub.20 aliphatic hydrocarbons containing a minor amount of cyclic 
components. The low temperature, high pressure distillate mode secondary 
stage operation favors the formation of linear or slightly-branched 
oligomers. 
HIGH SEVERITY REACTOR OPERATION 
Referring to FIG. 5 of the drawing, olefinic feedstock is supplied to the 
plant through fluid conduit 301 under steady stream conditions. This 
C.sub.2.sup.+ feedstream is pressurized by compressor 302 and then 
sequentially heated by passing through process heat exchange unit 304 and 
furnace 305 to achieve the temperature for catalytic conversion in reactor 
system 310, including plural reactor vessels 311A, B, C, etc. 
The reactor sub-system section shown consists of three downflow fixed bed, 
series reactors on line with heat exchanger cooling means 312 A, B, C, D, 
E between reactors and following the subsystem. The reactor configuration 
allows for any reactor to be in any position, A, B or C. The reactor in 
position A has the most aged catalyst and the reactor in position C has 
freshly regenerated catalyst. The cooled reactor effluent is first 
separated in a phase separator unit 315 is to provide a condensed 
C.sub.5.sup.+ hydrocarbon liquid stream 316 and an ethene-rich vapor 
stream 317 comprising C.sub.2 -C.sub.4 aliphatic hydrocarbons, along with 
any other unreacted gaseous components which might be present in the 
feedstock, such as hydrogen, carbon oxides, methane, nitrogen or other 
inert gases. Extraneous water may be removed from the system through 
separator line 318. 
Condensed hydrocarbon reactor effluent 316 separated from the effluent 
vapor is further fractionated. A stripping unit 320 which may be reboiled 
by exchanging a reactor effluent stream with tower bottoms in reboiler 
321, removes a significant fraction of dissolved light gases, including a 
minor amount of unreacted ethene. The C.sub.2.sup.- stripped gases are 
passed through conduit 322 operatively connecting the stripper with a 
downstream sorption unit 330. Ethane and heavier hydrocarbons are removed 
from the recycle loop through stripper 320. This tower may be designed to 
lose as little ethylene as possible while maintaining a reasonable tower 
bottom temperature. High pressure favors the split between ethylene and 
ethane. Preferably the liquid stripper effluent 324 is debutanized in a 
fractionation subsystem 340 to provide a C.sub.4.sup.- overhead stream, 
which is deethanized to provide LPG (C.sub.3 -C.sub.4 alkane) product 341 
and light offgas. The C.sub.5.sup.+ debutanizer bottom stream is split in 
an atmospheric distillation tower to provide raw distillate product stream 
342 and an olefinic gasoline stream 344 for recycle and/or recovery of a 
minor amount as raw gasoline product. 
To recycle unconverted ethylene, recycle gasoline is used to selectively 
absorb it in the ethylene absorber 330. Ethylene is recovered from the 
vapor stream 317 leaving the reactor effluent separator and from the 
stripper overhead 322. The H.sub.2, CO, CO.sub.2 and CH.sub.4 inerts which 
may enter with the feed are removed in the tower overhead via conduit 331 
to prevent their build up in the system. 
The gasoline sorbent is an aliphatic hydrocarbon mixture boiling in the 
normal gasoline range of about 50 to 165.degree. C. (125 to 330.degree. 
F.), with minor amounts of C.sub.4 -C.sub.5 alkanes and alkenes. 
Preferably, the total gasoline sorbent stream to ethylene sorbate mole 
ratio is greater than about 4:1. The process may be operated with a mole 
ratio of about 0.2 moles to about 10 moles of gasoline per mole of 
C.sub.2.sup.+ olefins in the feedstock. 
The tower pressure and bottom temperature may be selected such that enough 
CO.sub.2 leaves the system without carrying too much ethylene with it. 
Ethylene absorption efficiency can be improved if CO.sub.2 is removed by 
an optional amine scrubber or the like (not shown) before entering the 
tower. 
There is no need for a recycle compressor because all the recovered 
ethylene is dissolved in the recycle gasoline as a sorbate stream 332 and 
passed by pump 334 to the reactor. Advantageously, the liquid recycle 
stream is brought to process pressure before being heated to vaporize at 
least a portion of the olefinic components. 
It is understood that the various process conditions are given for a 
continuous system operating at steady state, and that substantial 
variations in the process are possible within the inventive concept. In 
the detailed examples, metric units and parts by weight are employed 
unless otherwise specified. 
The fractionation towers depicted in the drawing may employ a plate column 
in the primary tower and a packed column in the secondary tower, however, 
the fractionation equipment may also employ vapor-liquid contact means of 
various designs in each stage including packed beds of Raschig rings, 
saddles or other porous solids or low pressure drop contact devices 
(Glitsch grids). The number of theoretical stages will be determined by 
the feedstream composition, liquid:vapor (L/V) ratios, desired recovery 
and product purity. 
A typical high severity multi-zone reactor system employs inter-zone 
cooling, whereby the reaction exotherm can be carefully controlled to 
prevent excessive temperature above the normal moderate range of about 
260.degree. to 370.degree. C. 
Advantageously, the maximum temperature differential across any one reactor 
is about 30.degree. C. (.DELTA.T.about.50.degree. F.) and the space 
velocity (LHSV based on olefin feed) is about 0.5 to 1. Heat exchangers 
provide inter-reactor cooling and reduce the effluent to fractionation 
temperature. It is an important aspect of energy conservation in the MOGD 
system to utilize at least a portion of the reactor exotherm heat value by 
exchanging hot reactor effluent from one or more reactors with a 
fractionator stream to vaporize a liquid hydrocarbon distillation tower 
stream, such as the debutanizer bottoms. Optional heat exchangers may 
recover heat from the effluent stream prior to fractionation. Gasoline 
from the recycle conduit is pressurized by pump means and combined with 
feedstock, preferably at a mole ratio of about 2-3 moles per mole of 
olefin in the feedstock. It is preferred to operate the high severity 
reactors at elevated pressure of about 4200 to 7000 kPa (600-1000 psig), 
with a minimum olefin partial pressure of 1200 kPa at the reactor system 
inlet. 
The reactor system contains multiple downflow adiabatic catalytic zones in 
each reactor vessel. The liquid hourly space velocity (based on total 
fresh feedstock) is about 1 LHSV. In this mode the molar recycle ratio for 
gasoline is at least equimolar, based on total olefins in the fresh 
feedstock and recycle. The preferred molar ratio olefinic gasoline to 
fresh feedstock olefin is at least 2:1. This will also assure adequate 
sorbent for the portion unit. Typical reactor conditions are set forth in 
the following tables. 
TABLE IV 
______________________________________ 
HIGH SEVERITY REACTOR SYSTEM 
______________________________________ 
Feedstock Yield on Olefin Converted 
Component Wt % Component Wt % 
______________________________________ 
Inerts 5.00 CH.sub.4 0.10 
CH.sub.4 2.00 C.sub.2 H.sub.6 
3.90 
C.sub.2 H.sub.4 
81.20 C.sub.3 H.sub.8 
4.00 
C.sub.2 H.sub.6 
0.62 IC.sub.4 H.sub.10 
2.00 
C.sub.3 H.sub.6 
3.71 NC.sub.4 H.sub.10 
2.00 
C.sub.3 H.sub.8 
0.20 IC.sub.5 H.sub.12 
1.32 
IC.sub.4 H.sub.10 
0.25 NC.sub.5 H.sub.12 
0.09 
NC.sub.4 H.sub.10 
0.45 C.sub.5 H.sub.10 
2.99 
C.sub.4 H.sub.8 
0.12 C.sub.6 -300.degree. Gaso. 
39.60 
IC.sub.5 H.sub.12 
2.31 330.degree. + Dist. 
44.00 
NC.sub.5 H.sub.12 
0.10 
C.sub.5 H.sub.10 
1.73 
C.sub.6.spsb.+ 
2.31 
______________________________________ 
Conversion on Feed to Reactor 
Olefins 
Wt. % 
______________________________________ 
C.sub.2 
75 
C.sub.3 
95 
C.sub.4 
85 
______________________________________ 
TABLE V 
______________________________________ 
REACTOR CONDITIONS 
______________________________________ 
Space Velocity, LHSV 0.5 
(Based on olefins fed to reactor) 
Reactor A inlet pressure, psig 
900 
Minimum Olefin pp at reactor inlet, psia 
180 
Exothermic Heat of Reaction 
1040 
BTU/# olefins converted 
Rate of Heat Release Uniformly over bed 
Maximum Allowable 50 
.DELTA.T in Reactor, .degree.F. 
Reactor Inlet Temperature 
500/700.degree. F. 
SOC/EOC 
Gasoline Recycle, Mol/Mol Olefin Feed 
2:1 
Coke on Catalyst, wt. % SOC 
0 
EOC 30 
Cycle Length, Days 30 
Catalyst HZSM-5 
1/16" Extrudate 
______________________________________ 
More than 90% of ethylene is recovered in the above example from the 
effluent. 
While the invention has been described by specific examples and 
embodiments, there is no intent to limit the inventive concept except as 
set forth in the following claims.