Integrated process for converting methanol to gasoline and distillates

An integrated process comprising the steps of pressurizing a C.sub.3.sup.+ olefin hydrocarbon stream and a methanol feed and contacting the C.sub.3.sup.+ hydrocarbon stream and methanol feed in a first reaction zone with a medium-pore shape selective oligomerization zeolite catalyst at elevated pressure and moderate temperature to convert at least a portion of the C.sub.3.sup.+ hydrocarbons and methanol feed to heavier liquid hydrocarbon product stream comprising olefinic gasoline and distillate range liquids. An ethene stream and a stream containing unreacted methanol and water are recovered from the first reaction zone effluent and the methanol is separated from the water. The ethene and the separated unreacted methanol are contacted with a medium-pore shape selective zeolite catalyst in a second reaction zone at elevated temperature and moderate pressure to convert the methanol feed to hydrocarbons comprising C.sub.3.sup.+ olefins and cooling effluent from the second reaction zone to recover a C.sub.3.sup.+ olefin hydrocarbon stream and a C.sub.2.sup.- light gas stream. The C.sub.3.sup.+ olefin hydrocarbon stream is fed to the first reaction zone with the methanol feed.

BRIEF DESCRIPTION OF THE INVENTION 
The present invention relates to a process for the conversion of 
oxygenates, such as methanol or dimethyl ether (DME), to liquid 
hydrocarbons. The oxygenate feed is fed together with a light olefinic 
stream at elevated pressure and moderate temperature to a first reactor 
and contacted with an oligomerization zeolite catalyst to produce heavier 
liquid hydrocarbons. 
The present invention is particularly directed to a process for the 
conversion of methanol feed and olefin hydrocarbons to gasoline and 
distillate range liquid hydrocarbons at elevated pressure and moderate 
temperature. 
The present invention is more particularly directed to a process for the 
conversion of methanol and C.sub.3.sup.+ olefin hydrocarbons to gasoline 
and distillate range liquid hydrocarbons which comprises contacting the 
methanol and C.sub.3.sup.+ olefin hydrocarbons in a first reaction zone 
with a medium-pore shape selective oligomerization zeolite catalyst at 
elevated pressure and moderate temperature to convert at least a portion 
of the C.sub.3.sup.+ hydrocarbons and methanol feed to heavier liquid 
hydrocarbon product stream comprising olefinic gasoline and distillate 
range liquids. 
The present invention also directed to a MTO/MOGD process in which a 
methanol feed and olefin hydrocarbons are fed to the MOGD reactor at 
elevated pressure and moderate temperature to convert the methanol feed 
and olefin hydrocarbon feed to olefinic gasoline and distillate range 
liquids. 
The methanol feed that is not converted to olefins in the MOGD reactor, 
i.e. the unreacted methanol, is extracted from the MOGD effluent by water 
wash, optionally dewatered and then fed to the MTO reactor. An ethene 
containing stream is separated from the MOGD reactor effluent and can be 
optionally fed to the MTO reactor with the separated unreacted methanol at 
elevated temperature and moderate pressure to convert the ethene and 
separated unreacted methanol to C.sub.3.sup.+ olefin hydrocarbons. The 
C.sub.3.sup.+ olefin hydrocarbons are fed to the MOGD reactor with the 
original methanol feed. 
BACKGROUND 
In order to provide an adequate supply of liquid hydrocarbons for use as 
synfuels or chemical feedstocks, various processes have been developed for 
converting coal and natural gas to gasoline, distillate and lubricants. A 
substantial body of technology has grown to provide oxygenated 
intermediates, especially methanol. Large scale plants can convert 
methanol or similar aliphatic oxygenates to liquid fuels, especially 
gasoline. However, the demand for heavier hydrocarbons has led to the 
development of processes for increasing the yield of gasoline and diesel 
fuel by multi-stage techniques. 
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 Plank, Rosinski and Givens U.S. Pat. No(s). 3,960,978, and 4,021,502, 
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. No(s). 
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 ZSM-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 is 
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. Operating details for typical MOGD units are 
disclosed in Owen et al U.S. Pat. No(s). 4,445,031 and 4,456,779, and 
Tabak U.S. Pat. No. 4,433,185, 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 or corresponding ethers to olefins. 
Particular interest has been directed to a catalytic process ("MTO") for 
converting low cost methanol to valuable hydrocarbons rich in ethene and 
C.sub.3.sup.+ alkenes. Various processes are described in Batter et al 
U.S. Pat. No. 3,894,107, Chang et al U.S. Pat. No. 3,928,483, Lago U.S. 
Pat. No. 4,025,571, Daviduk et al U.S. Pat. No. 4,423,274, and Young U.S. 
Pat. No. 4,433,189, incorporated herein by "reference. It is generally 
known that the MTO process can be optimized to produce a major fraction of 
C.sub.2 -C.sub.4 olefins. "Prior process proposals have included a 
separation section to recover ethene and other gases from by-product water 
and C.sub.5.sup.+ hydrocarbon liquids. The oligomerization process 
conditions which favor the production of C.sub.10 -C.sub.20 and higher 
aliphatics tend to convert only a small portion of ethene as compared to 
C.sub.3.sup.+ olefins. 
The Gould et al U.S. Pat. No. 4,579,999 discloses an integrated process for 
the conversion of methanol to gasoline and distillate. In a primary 
catalytic stage (MTO) methanol is contacted with zeolite catalyst to 
produce C.sub.2 -C.sub.4 olefins and C.sub.5.sup.+ hydrocarbons. In a 
secondary catalytic stage (MOGD) containing an oligomerization catalyst 
comprising medium pore shape selective acidic zeolite at increased 
pressure, a C.sub.3.sup.+ olefins stream from the primary stage is 
converted to gasoline and/or distillate liquids. 
The Harandi et al U.S. Pat. No. 4,899,002 discloses a process for the 
increased production of olefinic gasoline, which comprises the integration 
of oxygenates to olefin (MTO) conversion with olefin to gasoline 
conversion under moderate severity conditions. The product of the olefins 
to gasoline conversion is passed to an olefin to gasoline and distillate 
(MOGD) conversion zone for distillate production. 
The methanol to olefin process (MTO) operates at high temperature and 
moderate pressure and high catalyst contact time in order to obtain 
efficient conversion of the methanol to olefins. These process conditions, 
however, produce an undesirable amount of aromatics and C.sub.2.sup.- 
light gas stream and require a large investment in plant equipment. 
The olefins to gasoline and distillate process (MOGD) operates at moderate 
temperatures and elevated pressures to produce olefinic gasoline and 
distillate products. When the conventional MTO process effluent is used as 
a feed to the MOGD process, the aromatic hydrocarbons produced in the MTO 
unit are desirably separated and a relatively large volume of MTO product 
effluent has to be cooled and treated to separate a C.sub.2.sup.- light 
gas stream, which is unreactive, except for ethene which is reactive to 
only a small degree, in the MOGD reactor, and the remaining hydrocarbon 
stream has to be pressurized to the substantially higher pressure used in 
the MOGD reactor. 
The problems to be solved were to reduce the overall size and investment in 
the MTO reactor, reduce the amount of the methanol feed fed to the MTO 
reactor in order that the process could be carried out under lower 
severity operating conditions which improves selectivity to not produce 
aromatics and not produce large amounts of C.sub.2.sup.- light gas. At the 
same time it was desired to maintain the total effective amount of the 
methanol feed converted to olefins and to improve the overall selectivity 
of the MTO/MOGD process to produce more olefinic gasoline and distillates. 
OBJECTS OF THE INVENTION 
It is an object of the present invention to improve the overall operation 
and cost of conversion of methanol to gasoline and distillate by a process 
which comprises feeding the entire methanol feed to the MOGD reactor 
together with a light olefin hydrocarbon stream. 
It is another object of the present invention to separate and recover 
unconverted methanol from the MOGD reactor effluent and to recycle a 
portion of the unconverted methanol to the MOGD reactor. 
It is another object of the present invention to improve the overall 
operation and cost of conversion of methanol to gasoline and distillate by 
process integration of a methanol to olefin conversion process with an 
olefin to gasoline and distillate conversion process. 
It is another object of the present invention to reduce the size and 
investment in the methanol to olefin conversion process by feeding the 
entire methanol feed to a methanol and olefin to gasoline or distillate 
conversion step (MOGD) and separating unreacted methanol and ethene, and 
feeding the unreacted methanol and ethene to the methanol to olefin 
conversion process (MTO). 
SUMMARY OF THE INVENTION 
In accordance with the present invention methanol, dimethyl ether (DME) or 
other lower oxygenates containing less than four carbon atoms may be 
converted to liquid fuels particularly gasoline and distillate, in a 
continuous process with integration between major process units. The 
methanol feed, together with a light olefin hydrocarbon stream, e.g. a 
C.sub.3.sup.+ olefin hydrocarbon stream, is fed to an olefin to gasoline 
and distillate unit reactor (MOGD) to produce gasoline and distillate. 
Unreacted methanol is recovered from the MOGD reactor effluent and is 
recycled to the MOGD reactor or is fed to a methanol to olefins reactor 
(MTO) to produce olefin hydrocarbon feed to the MOGD reactor. 
The present invention is specifically directed to an integrated process for 
the conversion of oxygenate feeds such as methanol and dimethyl ether to 
olefinic gasoline and distillate range liquid hydrocarbons. The process 
comprises the steps of pressurizing a C.sub.3.sup.+ olefin hydrocarbon 
stream and the methanol feed and contacting them in a first (MOGD) reactor 
with oligomerization catalyst at elevated pressure and moderate 
temperature to convert the C.sub.3.sup.+ olefin stream and methanol to a 
heavier liquid hydrocarbon stream comprising olefinic gasoline and 
distillate range hydrocarbons. The heavier liquid stream is cooled to 
preliminarily separate C.sub.3.sup.+ liquid hydrocarbons from 
C.sub.2.sup.- light gas, unreacted methanol and by-product water. The 
C.sub.3.sup.+ liquid hydrocarbons can be further treated to recover an LPG 
stream, a C.sub.5 -C.sub.9 olefinic gasoline stream and a C.sub.10 
-C.sub.20 distillate stream. The C.sub.10 -C.sub.20 distillate stream can 
be hydrotreated to produce high quality distillate product. 
The unreacted methanol is separated from the by-product water. The 
unreacted methanol, preferably along with the C.sub.2.sup.- light gas, is 
contacted with zeolite catalyst in a second reactor (MTO) at elevated 
temperature and moderate pressure to convert the unreacted methanol to 
C.sub.3.sup.+ light olefin hydrocarbons. The effluent from the second 
reactor is cooled to separate a C.sub.3.sup.+ olefin hydrocarbon stream 
and a C.sub.2.sup.- light gas stream. The C.sub.3.sup.+ hydrocarbon stream 
is fed with the original methanol feed to the first reactor as discussed 
above. 
Advantageously, the first and second reactors can contain ZSM-5 type 
zeolite catalyst.

DETAILED DESCRIPTION OF THE INVENTION 
Oxygenated Feed 
Numerous oxygenate organic compounds can be used as the feed to be 
converted to olefinic gasoline and distillate in the present invention. 
Since methanol or its ether derivative (DME) are industrial commodities 
from synthesis gas or the like processes, these materials are utilized in 
the description herein as preferred starting materials. It is understood 
by those skilled in the art that the methanol to olefin type processes can 
employ methanol, dimethyl ether and mixtures thereof, as well as other 
lower aliphatic alcohols and ethers, lower ketones and/or aldehydes. It is 
also understood by those skilled in the art to partially convert 
oxygenates, such as methanol, by dehydration, as in the catalytic reaction 
of methanol over gamma-alumina to produce DME intermediate. Typically, an 
equilibrium mixture (CH.sub.3 OH.revreaction.CH.sub.3 OCH.sub.3 +H.sub.2 
O) is produced by partial dehydration. In the first reactor, methanol and 
olefins are converted to gasoline and distillate (MOGD); and in the second 
reactor, unreacted methanol from the first reactor and olefins are 
converted to lower olefins (MTO). 
Catalyst 
Catalyst versatility permits the same zeolite catalyst to be used in the 
first reactor unit oligomerization stage (MOGD) and in the second reactor 
unit methanol to olefins stage (MTO). While it is within the inventive 
concept to employ substantially different catalysts in these reactors, it 
is advantageous to employ a standard ZSM-5 catalyst having a silica to 
alumina molar ratio of 70:1 in the first and second reactors. 
Recent developments in zeolite technology have provided a group of 
medium-pore shape-selective siliceous materials having similar pore 
geometry. Most prominent among these intermediate pore size zeolites is 
ZSM-5, which is usually synthesized with Bronsted acid active sites by 
incorporating a tetrahedrally coordinated metal, such as Al, Ga, B or Fe, 
within the zeolitic framework. These medium-pore zeolites are favored for 
acid catalysis; however, the advantages of ZSM-5 structures may be 
utilized by employing highly siliceous material or crystalline 
metallosilicate having one or more tetrahedral species having varying 
degrees of acidity. ZSM-5 crystalline structure is readily recognized by 
its X-ray diffraction pattern, which is described in Argauer et al U.S. 
Pat. No. 3,702,866, incorporated by reference. 
The zeolite catalysts preferred for use herein include the medium-pore 
(i.e., about 5-7 A) 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 1-200. In an operating 
reactor the coked catalyst may have an apparent activity (alpha value) of 
about 1 to 80 under the process conditions to achieve the required degree 
of reaction severity. Representative of the ZSM-5 zeolites are ZSM-5, 
ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and MCM-22. ZSM-5 
is disclosed in U.S. Pat. No. 3,702,886 and U.S. Pat. Re. No. 29,948. The 
ZSM-5, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48 and MCM-22 catalyst are 
preferred. The MCM-22 catalyst is described in U.S. Pat. No. 4,954,325. 
Other suitable zeolites are disclosed in U.S. Pat. No(s). 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, 4,417,086, 4,517,396, and 4,542,251. The disclosures of the 
above mentioned patents are incorporated herein by reference. While 
suitable zeolites having a coordinated silica to metal oxide molar ratio 
of 20:1 to 200:1 or higher may be used, it is advantageous to employ a 
standard ZSM-5 having a silica to alumina molar ratio of about 25:1 to 
70:1. A typical zeolite catalyst component having Bronsted acid sites may 
consist essentially of aluminosilicate ZSM-5 zeolite with 5 to 95 wt. % 
silica and/or alumina binder. 
Certain of the ZSM-5 type medium pore shape-selective catalysts are 
sometimes known as pentasils. In addition to the preferred 
aluminosilicates, the borosilicate, ferrosilicate and "silicalite" 
materials may be employed. It is advantageous to employ a standard ZSM-5 
having a silica:alumina molar ratio of 25:1 to 70:1 with an apparent alpha 
value of 1-80. 
ZSM-5 type pentasil zeolites are particularly useful in the process because 
of their regenerability, long life and stability under the extreme 
conditions of operation. Usually the zeolite crystals have a crystal size 
from about 0.01 to over 2 microns or more, with 0.02-1 micron being 
preferred. The zeolite catalyst crystals are normally bound with a 
suitable inorganic oxide, such as silica, alumina, etc. to provide a 
zeolite concentration of about 5 to 95 wt. %. A preferred catalyst 
comprises 25% to 65% H-ZSM-5 catalyst contained within a silica-alumina 
matrix binder and having a fresh alpha value of less than 600. 
When employing a ZSM-5 type zeolite catalyst in a fluidized bed as a fine 
powder such a catalyst should comprise the zeolite suitably bound or 
impregnated on a suitable support with a solid density (weight of a 
representative individual particle divided by its apparent "outside" 
volume) in the range from 0.6-2 g/cc, preferably 0.9-0.6 g/cc. The 
catalyst particles can be in a wide range of particle sizes up to about 
250 microns, with an average particle size between 20 and 100 microns, 
preferabaly in the range of 10-150 microns and with the average particle 
size between 40 and 80 microns. When these solid particles are placed in a 
reactor bed where the superficial fluid velocity is 0.3-2 ft./sec., 
fluidized bed operation is obtained. The velocity specified here is for an 
operation at a total reactor pressure of about 0 to 30 psig (100 to 300 
kPa). Those skilled in the art will appreciate that at higher pressures, a 
lower gas velocity may be employed to ensure proper fluidized bed 
operation. 
In the fluidized bed embodiments of the present invention it is 
advantageous to employ a particle size range consisting essentially of 1 
to 150 microns. Average particle size is usually about 20 to 100 microns, 
preferably 40 to 80 microns. Particle distribution may be enhanced by 
having a mixture of larger and smaller particles within the operative 
range, and it is particularly desirable to have a significant amount of 
fines. Close control of distribution can be maintained to keep about 10 to 
25 wt. % of the total catalyst in the reaction zone in the size range less 
than 32 microns. This class of fluidizable particles is classified as 
Geldart Group A. 
The light olefin production is promoted by the zeolite catalysts having a 
high concentration of Bronsted acid reaction sites. Accordingly, an 
important criterion is selecting and maintaining the catalyst to provide 
either fresh catalyst having acid activity or by controlling catalyst 
deactivation and regeneration rates to provide an apparent average alpha 
value of about 1 to 50, preferably 1 to 10. 
A further description of the zeolite catalyst is found in Owen et al U.S. 
Pat. No. 4,456,779 which is incorporated herein by reference. 
The oligomerization catalysts preferred for use herein in the MOGD fixed 
bed reactor include the 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. A suitable catalyst for 
fixed bed operation is 65 wt. % HZSM-5 zeolite with an alumina binder in 
the form of cylindrical extrudates of about 1-5 mm. Other catalysts and 
processes suitable for converting methanol/DME to lower olefins are 
disclosed in Bonafaz U.S. Pat. No. 4,393,265, Vogt et al U.S. Pat. No. 
4,387,263, and Marosi et al European Patent Application No. 0081683. 
The ZSM-5 type catalysts are particularly advantageous for use in the 
present invention because the same material may be employed for 
dehydration of methanol to DME, conversion of methanol to lower olefins 
and oligomerization reactions. A particular advantage of the process of 
the present invention is that the spent catalyst from a higher pressure 
unit can be used in a lower pressure unit as fresh make-up catalyst. 
REACTORS 
A number of different types of reactors can be used in the above mentioned 
processes including tubular, moving bed, fixed bed, fluid bed and riser 
reactors. The preferred reactor types and the operations of each in the 
processes are briefly described below. 
MOGD Reactor I 
The MOGD reaction is preferably carried out in a fixed bed multi-stage 
reactor. Suitable reactor designs, process conditions and techniques are 
described in Harandi et al U.S. Pat. No(s). 4,777,316, and 4,877,921, 
Gould U.S. Pat. No. 4,579,999, and Owen et al U.S. Pat. No. 4,456,779, all 
of which are incorporated herein by reference thereto. 
Though a fixed bed multi-stage reactor is preferred for the MOGD reaction 
system, in certain circumstances a fluidized bed reactor system can be 
used. For example, see Harandi et al U.S. Pat. No. 4,877,921 which is 
incorporated herein by reference thereto. 
MTO Reactor II 
The MTO reaction is preferably carried out in a fluidized bed reactor 
because of the highly exothermic nature of the methanol to olefin 
reaction. In a preferred MTO reactor, a bed of finely divided (&lt;150 
microns) ZSM-5 catalyst is maintained in a turbulent fluidization regime. 
Hot feedstock vapor is passed upwardly through the fluidized bed at a 
superficial velocity of about 0.3 to 2 meters per second, maintaining a 
bed density of about 200 to 600 kg/m.sup.3. By operating at about 
520.degree. C..+-.20.degree. C. and a catalyst activity sufficient to 
yield a propane:propene ratio of about 0.02 to 0.3:1, the production of 
ethylene and C.sub.3.sup.- paraffins can be controlled at a low level. 
A suitable reactor and operating technique for carrying out this step of 
the invention are disclosed in Gould et al Ser. No. 687,045, filed Dec. 
28, 1984. Other fluidized bed reactor systems suitable for use in the MTO 
reactor stage are disclosed in Avidan U.S. Pat. No. 4,746,762 and Harandi 
et al U.S. Pat. No. 4,777,316, both of which are incorporated herein by 
reference. 
Though fluidized bed reactors are preferred for the MTO reactor stage, 
multi-stage fixed bed reactors provided with inter-stage cooling can also 
be used. See, for example, Graziani et al U.S. Pat. No. 4,542,252. 
In the description used in the present application, metric units and parts 
by weight are employed unless otherwise stated. Various reactor 
configurations ,au be used, including fluidized bed catalytic reactors, 
moving bed and multi-stage fixed bed reactors. 
The invention is further described with reference to the figures of the 
drawings. 
DESCRIPTION OF THE FIGURES OF THE DRAWINGS 
Referring to FIG. 1 methanol feed in line 1 is fed to the MOGD reactor I to 
selectively produce distillate hydrocarbons. The methanol feed can contain 
0 to 50% by weight water, generally 5 to 40% by weight water, typically 4 
to 20% by weight water. Preferably the methanol feed contains less than 5% 
by weight water. A C.sub.3.sup.+ olefin hydrocarbon feed, which is 
typically derived from a thermal or catalytic cracking operation, is fed 
through line 4 to the MOGD reactor I. 
The methanol feed fed through line 1 and C.sub.3.sup.+ olefin feed fed 
through line 4 are contacted in reactor I with HZSM-5 catalyst where the 
methanol is converted to lower olefins, gasoline, distillate, and 
by-product water. The water is formed by the dehydration of the methanol. 
The C.sub.3.sup.+ olefin hydrocarbon feed fed through line 4 can have the 
following composition. 
______________________________________ 
Wt. % Typically Wt. % 
______________________________________ 
C.sub.2.sup.- 
0-14 0-5 
C.sub.3 -C.sub.4 
0-100 90-99 
C.sub.5 -C.sub.9 
100-0 0-5 
C.sub.10 -C.sub.20 
0-5 0 
______________________________________ 
The C.sub.2.sup.- stream and the C.sub.10 -C.sub.20 stream are normally 
removed prior to feeding the stream to the MOGD reactor. 
A suitable C.sub.3.sup.+ olefin feed stream can be obtained from a 
dehydrogenation plant, e.g. dehydrogenation of a propane stream. 
The methanol feed and the remaining C.sub.3.sup.+ olefin feed are fed to 
reactor I at 0.01 to 10 WHSV, preferably 0.1 to 5.0 WHSV and more 
preferably at 0.3 to 1.0 WHSV, based on olefins and methanol content of 
the feed. The mixed methanol and C.sub.3.sup.+ olefin hydrocarbon feed is 
contacted in reactor I with an HZSM-5 catalyst arranged in a multi-stage 
fixed bed reactor preferably having three to four stages. Prior to 
entering the reactor I the mixed feed is pressurized by means not shown to 
the elevated pressure of the reactor I. 
The reactor I is operated to selectively produce distillate hydrocarbons at 
a temperature of 177 to 371.degree. C. (350 to 700.degree. F.), preferably 
204 to 343.degree. C. (400 to 65.degree. F.) and more preferably 204 to 
316.degree. C. (400 to 600.degree. F.). The reactor I is operated at a 
pressure of 4237 to 20,780 kPa (600 to 3000 psig), preferably 4237 to 
10,440 kPa (600 to 1500 psig) and more preferably 5600 to 7000 kPa (800 to 
1000 psig). 
After separation of any unconverted methanol and water present, the reactor 
I hydrocarbon effluent has the following composition. 
______________________________________ 
Typically Wt. % 
______________________________________ 
C.sub.2.sup.- 
1-2 
C.sub.3 -C.sub.4 
8-12 
C.sub.5 -C.sub.9 
15-30 
C.sub.10 -C.sub.20 
60-75 
______________________________________ 
The aromatics content of gasoline fraction is 0-12 wt. %, based on total 
converted hydrocarbons. 
At least 50% and preferably at least 95% of the methanol feed is converted 
to hydrocarbons. 
The reactor I effluent is separated by means not shown into a C.sub.2.sup.- 
light gas stream withdrawn through line 10; a gas stream comprising 
C.sub.3 -C.sub.4 (LPG) hydrocarbons withdrawn through line 7; a C.sub.5 
-C.sub.9 olefinic gasoline stream withdrawn through line 8; and a C.sub.10 
-C.sub.20 distillate hydrocarbon stream Withdrawn through line 9. The 
unreacted methanol and water present are separated by cooling, phase 
separation, and in some cases by water washing of the hydrocarbon effluent 
leaving reactor I, and are withdrawn through line 11 and are fed to 
methanol-water separator 12, which is preferably a distillation tower. The 
methanol is separated from the water. The water is withdrawn through line 
13. The unreacted methanol is withdrawn through line 14. At least a 
portion of the unreacted methanol can be recycled through line 15 to the 
methanol line 1. The C.sub.2.sup.- light gas comprises mostly ethane and 
ethene. It may also contain small amounts of methane and hydrogen. 
In an embodiment of the present invention the methanol and water stream 
withdrawn from the MOGD reactor in line 11 is contacted with the 
C.sub.3.sup.+ olefin feed in line 4, by means not shown, to extract 
methanol from water. The C.sub.3.sup.+ olefin feed containing the 
extracted methanol is then fed to the MOGD reactor. This extraction step 
reduces the distillation requirement for the methanol-water separations. 
The higher pressures increase selectivity to distillate hydrocarbons. At 
any particular selected pressure the lower temperatures increase 
selectivity to distillate. 
Referring to FIG. 2, the methanol feed is fed through line 1 to the MOGD 
reactor I. The methanol feed can contain 0 to 50% by weight water, 
generally 5 to 40% by weight water, typically 4 to 20% by weight water. 
Preferably the methanol feed contains less than 5% by weight water. A 
C.sub.3.sup.+ olefin hydrocarbon feed is fed through line 4 to the MOGD 
reactor I. The C.sub.3.sup.+ olefin feed stream can be supplemented with 
C.sub.3.sup.+ hydrocarbons derived from a thermal or catalytic cracking 
operation. The process is described with reference to producing gasoline 
and distillate boiling range hydrocarbons. 
The C.sub.3.sup.+ olefin containing stream and the methanol feed are fed 
through lines 4 and 1, respectively, to the MOGD reactor I. The 
C.sub.3.sup.+ olefin feed and the methanol can be mixed before entering 
the reactor or after they enter the reactor. The methanol and the 
C.sub.3.sup.+ olefin are fed to reactor I at 0.01 to 10 WHSV, preferably 
0.1 to 5.0 WHSV and more preferably 0.3 to 1.0 WHSV, based on olefins plus 
methanol. The mixed methanol and C.sub.3.sup.+ olefin hydrocarbon feed is 
contacted in reactor I with an HZSM-5 catalyst arranged in a multi-stage 
fixed bed reactor preferably having three or four stages. Prior to 
entering the reactor I the feed is pressurized by means not shown to the 
substantially elevated pressure of the reactor I. 
The reactor I is operated at a temperature of 177 to 371.degree. C. (350 to 
700.degree. F.), preferably 204 to 343.degree. C. (400 to 650.degree. F.) 
and more preferably 204 to 316.degree. C. (400 to 600.degree. F.) The 
reactor I is operated at a pressure of 5237 to 20,780 kPa (600 to 3000 
psig), preferably 4237 to 10,440 kPa (600 to 1500 psig) and more 
preferably 5600 to 7000 kPa (800 to 1000 psig). At the lower pressures 
gasoline products are selectively produced, while at the higher pressures 
lubricating oil stocks are selectively produced. 
After separation of unconverted methanol and water, the reactor I 
hydrocarbon effluent has the following composition. 
______________________________________ 
Typically Wt. % 
______________________________________ 
C.sub.2.sup.- 
1-2 
C.sub.3 -C.sub.4 
8-12 
C.sub.5 -C.sub.9 
15-30 
C.sub.10 -C.sub.20 
60-75 
______________________________________ 
Aromatics content of gasoline fraction is 0-12 wt. %, based on total 
hydrocarbons. 
At least 50% and preferably at least 95% of the methanol feed is converted 
to olefin hydrocarbons. 
The reactor I effluent is separated by means not shown into a C.sub.2.sup.- 
light gas stream withdrawn through line 10 which is fed to reactor II or 
removed from the system through lines 10a and 5; a gas stream comprising 
C.sub.3 -C.sub.4 (LPG) hydrocarbons withdrawn through line 7; a C.sub.5 
-C.sub.9 olefinic gasoline stream withdrawn through line 8; and a C.sub.10 
-C.sub.20 distillate hydrocarbon stream withdrawn through line 9. The 
unreacted methanol and water are separated by cooling and phase separation 
and are withdrawn through line 11 and are fed to methanol-water separator 
12. The recovered unreacted methanol is separated from the water. The 
water is withdrawn through line 13 and line 6. The unreacted methanol is 
withdrawn through line 14 and fed to the MTO reactor II. Unreacted 
methanol may also be recovered by water washing the MOGD hydrocarbon 
effluent stream by means not shown and the methanol fed to the MTO reactor 
II. 
The separated unreacted methanol in line 14, preferably together with the 
separated C.sub.2.sup.- light gas stream withdrawn from the MOGD reactor 
I, is fed to reactor II and contacted with HZSM-5 catalyst. The methanol 
is converted to lower olefins and gasoline and by-product water. The 
olefin content of the C.sub.2.sup.- light gas can also be upgraded to 
heavier hydrocarbons in the MTO reactor II. 
The unreacted methanol in line 14 is fed to reactor II at 0.2 to 200 WHSV, 
preferably 0.3 to 3 WHSV and more preferably 0.5 to 2 WHSV. The reactor II 
is operated preferably as a dense fluidized bed at elevated temperatures 
of 260 to 538.degree. C. (500 to 1000.degree. F.), preferably 427 to 
510.degree. C. (800 to 950.degree. F.), and more preferably 471 to 
504.degree. C. (880 to 940.degree. F.). The reactor II is operated at 
moderate pressures of 101 to 789 kPa (0 to 100 psig), preferably 170 to 
652 kPa (10 to 80 psig), and more preferably 239 to 308 kPa (20 to 30 
psig). The reactor II effluent is cooled by means not shown and the 
by-product water is recovered by phase separation and withdrawn through 
line 6. After separation of water, the reactor II hydrocarbon effluent has 
the following composition. 
______________________________________ 
Typically Wt. % 
______________________________________ 
C.sub.2.sup.- 
10-14 
C.sub.3 -C.sub.4 
45-55 
C.sub.5 -C.sub.9 
25-35 
C.sub.10 -C.sub.20 
0-5 
______________________________________ 
Aromatics content of gasoline fraction is typically 10-14 wt. %, based on 
total hydrocarbons. 
At least 70% and preferably at least 99.9% of the methanol feed is 
converted to olefin hydrocarbons. 
The effluent hydrocarbons are treated to separate an overhead C.sub.2.sup.- 
light gas stream which is withdrawn through line 5 and a liquid 
C.sub.3.sup.+ olefin hydrocarbon stream which is withdrawn through line 4. 
The C.sub.3.sup.+ olefin hydrocarbon stream withdrawn through line 4 and 
the methanol feed in line 1 are fed to the MOGD reactor I as described 
above. 
Water and methanol present in the MOGD reactor effluent are separated and 
withdrawn through line 11 and fed to methanol-water separator 12. The 
methanol is separated from the water. The water is withdrawn through line 
13. The unreacted methanol is withdrawn though line 14. The unreacted 
methanol is fed through line 14 to MTO reactor II and the water is 
withdrawn through line 6. In the event there is any appreciable amount of 
unreacted methanol present in streams 7 and/or 8, the streams can be 
washed with water to eliminate methanol in these streams and the recovered 
methanol can be fed to reactor II. 
At the higher pressures a substantial amount of C.sub.20.sup.+ hydrocarbons 
are obtained which can be hydrotreated and used as lubricant stock. The 
relative proportion of C.sub.5 -C.sub.9 gasoline and C.sub.10 -C.sub.20 
distillate is determined by the reaction conditions in reactor I and the 
recycle rate of the C.sub.5 -C.sub.9 hydrocarbon fraction. The higher 
temperatures and lower pressures favor the C.sub.5 -C.sub.9 gasoline 
production and the lower temperatures, higher pressures and C.sub.5 
-C.sub.9 gasoline recycle favor heavy C.sub.10 -C.sub.20 distillate 
production. The C.sub.5 -C.sub.9 gasoline fraction withdrawn through line 
8 can optionally be recycled by means not shown to the MOGD reactor I to 
increase the production of distillate. 
The lower pressures increase the selectivity to gasoline hydrocarbons, 
while the higher pressures increase selectivity to distillate 
hydrocarbons. 
At any particular selected pressure the higher temperatures, for example 
371 to 538.degree. C. (700 to 1000.degree. F.), preferably 371 to 
482.degree. C. (700 to 900.degree. F.) and more preferably 371 to 
427.degree. C. (700 to 800.degree. F.), increase selectivity to gasoline, 
while the lower temperatures 260 to 371.degree. C. (500 to 700.degree. 
F.), preferably 288 to 343.degree. C. (550 to 650.degree. F.), and more 
preferably 316 to 343.degree. C. (600 to 650.degree. F.), increase 
selectivity to distillate. 
The MOGD reactor I can be operated at a temperature of 177 to 538.degree. 
C. (350 to 1000.degree. F.), preferably 204 to 482.degree. C. (400 to 
800.degree. F.), and more preferably 204 to 371.degree. C. (400 to 
700.degree. F.). The MOGD reactor I can be operated at a pressure of 308 
to 20,780 kPa (30 to 3000 psig), preferably 308 to 10,440 kPa (30 to 1500 
psig), and more preferably 308 to 7000 kPa (30 to 1000 psig). 
MOGD Gasoline Mode 
When the process is carried out to selectively produce gasoline boiling 
range hydrocarbons the methanol feed, and the C.sub.3.sup.+ olefin 
hydrocarbon feed from MTO reactor II are fed to the MOGD reactor I at 0.01 
to 100 WHSV, preferably 0.1 to 5.0 WHSV and more preferably 0.3 to 1.0 
WHSV. The mixed methanol and C.sub.3.sup.+ olefin hydrocarbon feed is 
contacted in reactor I as before with an HZSM-5 catalyst arranged in a 
multi-stage fixed bed reactor. 
The reactor I is operated at a temperature of 260 to 583.degree. C. (500 to 
1000.degree. F.), preferably 343 to 482.degree. C. (650 to 800.degree. 
F.), and more preferably 343 to 371.degree. C. (650 to 700.degree. F.). 
The reactor I is operated at a pressure of 308 to 10,440 kPa (30 to 1500 
psig), preferably 308 to 5600 kPa (30 to 800 psig) and more preferably 308 
to 4237 kPa (30 to 600 psig). At the conditions recited gasoline boiling 
range hydrocarbon products are selectively produced. 
MOGD Distillate Mode 
When the process is carried out to selectively produce distillate boiling 
range hydrocarbons, the C.sub.3.sup.+ olefin hydrocarbon feed from the MTO 
reactor II and the methanol feed are fed to the MOGD reactor I at 0.01 to 
10 WHSV, preferably 0.1 to 5.0 WHSV and more preferably at 0.3 to 1.0 
WHSV. The mixed methanol and C.sub.3.sup.+ olefin hydrocarbon feed is 
contacted in reactor I with an HZSM-5 catalyst arranged in a multi-stage 
fixed bed reactor. 
The reactor I is operated at a temperature of 177 to 371.degree. C. (350 to 
700.degree. F.), preferably 204 to 343.degree. C. (400 to 650.degree. F.) 
and more preferably 204 to 316.degree. C. (400 to 600.degree. F.). The 
reactor I is operated at a pressure of 4237 to 20,780 kPa (600 to 3000 
psig), preferably 4237 to 10,440 kPa (600 to 1500 psig) and more 
preferably 5600 to 7000 kPa (800 to 1000 psig). At the conditions recited 
distillate range hydrocarbons are selectively produced. 
The present invention is illustrated by the following Example. 
EXAMPLE 
This Example is described with reference to FIG. 1 of the drawings. The 
FIG. 1 is a schematic flow sheet of the process of the present invention. 
In this embodiment the process is carried out to selectively produce 
distillate hydrocarbons. 
A C.sub.3.sup.+ olefin hydrocarbon stream having the following composition 
is fed through line 4. 
______________________________________ 
Wt. % 
______________________________________ 
C.sub.2.sup.- 
12 
C.sub.3 -C.sub.4 
52 
C.sub.5 -C.sub.9 
33 
C.sub.10 -C.sub.20 
3 
______________________________________ 
The aromatic content of the gasoline fraction is 12 wt. %, based on total 
hydrocarbons. 
Where the process is carried out to selectively produce distillate 
hydrocarbons, the C.sub.2.sup.- light gas stream, the aromatic 
hydrocarbons and the C.sub.10 -C.sub.20 hydrocarbons are preferably 
removed. 
The MOGD reactor I is operated in the distillate mode. The pressure of the 
C.sub.3.sup.+ hydrocarbons is increased by pump means not shown to 5600 to 
7000 kPa (800 to 1000 psig) and fed to separation means not shown. In the 
separation means an overhead C.sub.2.sup.- light gas stream, a C.sub.5 
-C.sub.9 aromatics stream and a C.sub.10 -C.sub.20 stream are removed. The 
remaining C.sub.3 -C.sub.4 olefin hydrocarbons and C.sub.5 -C.sub.9 
hydrocarbons are then fed through line 4 to the MOGD reactor I. The 
pressure of the methanol feed is increased to 5600 to 7000 kPa (800 to 1 
psig) by pump means not shown and fed through line 1 to the MOGD reactor 
I. The feed rate of the combined C.sub.3 -C.sub.4 and C.sub.5 -C.sub.9 
hydrocarbon stream in line 4 and he methanol stream in line 1 to the MOGD 
reactor I is 0.3 to 1.0 WHSV. The reactor I feed includes a 3:1 weight 
ratio of recycle of the line 8 C.sub.5 -C.sub.9 product stream by means 
not shown. 
The methanol and C.sub.3 -C.sub.4 and C.sub.5 -C.sub.9 olefin hydrocarbons 
feed and recycle C.sub.5 -C.sub.9 hydrocarbons are mixed in the reactor I 
and contacted with HZSM-5 catalyst in a multi-stage fixed bed reactor 
having three stages. The reactor I is operated under conditions to 
optimize distillate product at a pressure of 5600 to 7000 kPa (800 to 1000 
psig) and at a temperature of 204 to 316.degree. C. (400 to 600.degree. 
F.). A portion or all of the C.sub.5 -C.sub.9 fraction withdrawn through 
line 8 can be recycled to the MOGD reactor. 
The reactor I effluent hydrocarbon product is withdrawn and separated into 
the desired process streams. Water and unconverted methanol are separated 
from the hydrocarbons and are withdrawn through line 11. The hydrocarbon 
portion of the MOGD reactor effluent typically has the following 
composition. 
______________________________________ 
Wt. % 
______________________________________ 
C.sub.2.sup.- 
1.0 
C.sub.3 -C.sub.4 
7.0 
C.sub.5 -C.sub.9 
29 
C.sub.10 -C.sub.20 
63 
______________________________________ 
The methanol conversion to hydrocarbons is about 70%. 
The C.sub.2.sup.- light gas in line 10 can be removed for ethene recovery. 
The MOGD effluent hydrocarbon stream is fractionated in a conventional 
manner with the C.sub.3 C.sub.4 (LPG) being removed through line 7; the 
C.sub.5 -C.sub.9 olefin gasoline being removed through line 8; and the 
C.sub.10 -C.sub.20 distillate stream being removed through line 9. The 
unconverted methanol and water in line 11 are fed to separator 12 in which 
the methanol is separated from the water. The unconverted methanol can be 
recycled to the MOGD reactor. 
The MTO reactor in the conventional MTO/MOGD process operates at higher 
temperatures than applicant's MOGD reactor. The higher temperature 
operation of the MTO reactor results in the conversion of a portion of the 
methanol feed to aromatic hydrocarbons and to C.sub.2.sup.- light gas. In 
accordance with applicant's invention, in which all of the original 
methanol feed is fed to the MOGD reactor, the overall amounts of aromatic 
hydrocarbon and C.sub.2.sup.- light gas products are significantly 
reduced. 
Further, if applicant's entire methanol feed were fed to applicant's 
invention MTO reactor (FIG. 2), the size of MTO reactor and the investment 
needed for the MTO plant would be about three times as large. 
The prior art process of feeding all of the methanol feed to the MTO 
reactor results in the production of more C.sub.2.sup.- light gas and in 
the production of more aromatic hydrocarbons, both of which decrease the 
desired gasoline and heavy distillate production, respectively. In 
addition, an increased aromatics removal capacity is needed to remove the 
relatively larger amount of aromatics from the MTO effluent hydrocarbon 
product, since it is preferred not to have too large amount of aromatics 
the gasoline product and to not have any significant amount of aromatics 
in the heavy distillate product. 
The foregoing description of the present invention has omitted various 
heating and cooling apparatus, catalyst regenerators, compressors and like 
equipment which are conventional and well known to those skilled in the 
art. Further, recycle streams other than those described can be utilized 
to optimize specifically desired process streams. 
The described integrated processes provide effective means for converting 
oxygenated organic compounds such as methanol, DME, lower aliphatic 
ketones and aldehydes to valuable hydrocarbon products. Thermal 
integration is achieved by employing heating and cooling means between 
various process streams, towers, absorbers, etc., in a conventional 
manner. 
Various modifications can be made to the systems, especially in the choice 
of equipment and non-critical processing steps. While the invention has 
been described by specific examples, there is no intent to limit the 
inventive concept except as set forth in the following claims.