Multistage process for oxygenate conversion to hydrocarbons

A continuous multistage process for preparing gasoline and/or distillate range hydrocarbons from lower molecular weight oxygenate feedstock wherein hydrocarbon yield is increased by recovering a vapor stream rich in ethene from an oxygenates conversion stage and reacting the ethene in a high severity reaction zone containing high activity zeolite catalyst.

FIELD OF THE INVENTION 
This invention relates to a multi-stage process for converting organic 
oxygenate material, such as methanol or dimethylether (DME) to liquid 
hydrocarbons. In particular it provides a continuous process for producing 
gasoline and/or distillate range hydrocarbons by oligomerizing olefins 
obtained from a methanol-to-olefins ("MTO") type process. 
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 + 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. 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 + 
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. Operating details for typical MOGD units are disclosed in 
copending U.S. patent applications Ser. No. 488,834, filed Apr. 26, 198 
(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 
lower alkenes. Various processes are described in U.S. Pat. No. 3,894,107 
(Butter et al), U.S. Pat. No. 3,928,483 (Chang et al), U.S. Pat. No. 
4,025,571 (Lago), U.S. Pat. No. 4,506,106 (Hsia et al), and U.S. Pat. No. 
4,579,999 (Gould et al). Significance of the methanol-to-olefins ("MTO") 
type processes, especially for producing ethene, is discussed in 
Hydrocarbon Processing, Nov. 1982, pp. 117-120. 
SUMMARY OF THE INVENTION 
The present invention is a continuous multi-stage process for increasing 
the yield of liquid hydrocarbons boiling in the gasoline and distillate 
range from a feedstock comprising C.sub.1 -C.sub.4 aliphatic oxygenates. 
The process comprises contacting the feedstock in a primary fluidized bed 
reaction zone with a shape selective medium pore crystalline 
metallosilicate catalyst under oxygenate conversion conditions to obtain a 
primary effluent comprising ethylene and C.sub.3 + olefins; withdrawing 
the primary effluent; separating the primary effluent into a gaseous 
stream rich in ethylene and a liquid stream comprising C.sub.3 + olefins; 
contacting the liquid stream comprising C.sub.3 + olefins with a shape 
selective medium pore crystalline metallosilicate catalyst in a fixed bed 
reaction zone at elevated temperature and pressure to obtain a product 
stream comprising hydrocarbons boiling in the gasoline and distillate 
range; contacting the ethylene-rich stream with a shape selective medium 
pore crystalline metallosilicate catalyst in a secondary fluidized bed 
reaction zone under oligomerization conditions to obtain a secondary 
effluent comprising unreacted ethylene and heavier hydrocarbons; 
withdrawing the secondary effluent; separating the secondary effluent into 
a gaseous stream comprising ethylene and a liquid stream comprising 
heavier hydrocarbons; and adding at least a portion of the liquid stream 
comprising heavier hydrocarbons to the fixed bed reaction zone. A product 
comprising hydrocarbons boiling in the gasoline and/or distillate range is 
withdrawn from the fixed bed reaction zone. 
Other objects and features of the invention will be seen in the following 
description and drawings.

DETAILED DESCRIPTION OF THE INVENTION 
In this description, metric units and parts by weight are employed unless 
otherwise stated. 
Catalyst versatility permits similar zeolites to be used in the fluidized 
bed oxygenate conversion reactor, the fluidized bed olefins 
oligomerization reactor, and the fixed bed MOGD reactor. The 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 fresh acid cracking 
activity of about 10-250. 
Recent developments in zeolite technology have provided a group of medium 
pore 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 materials 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 U.S. Pat. No. 3,702,866 (Argauer, et al.), 
incorporated by reference. 
The oligomerization catalysts preferred for use herein include the medium 
pore (i.e., about 5-7.ANG.) shape-selective crystalline aluminosilicate 
zeolites having a silica-to-alumina ratio of at least 12, a constraint 
index of about 1 to 12 and fresh acid cracking alpha-value of about 
10-250. 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 (Argauer et al). Other suitable 
zeolites are disclosed in U.S. Pat. Nos. 3,709,979 (Chu); U.S. Pat. No. 
4,076,979 (Walter et al), U.S. Pat. No. 3,832,449 (Rosinski et al); U.S. 
Pat. No. 4,076,842 (Plank et al); U.S. Pat. No. 4,016,245 (Plank et al); 
U.S. Pat. No. 4,414,423 (Miller); U.S. Pat. No. 4,417,086 (Miller); U.S. 
Pat. No. 4,517,396 (Hoek et al) and U.S. Pat. No. 4,542,251 (Miller). The 
disclosures of these patents are incorporated herein by reference. While 
suitable zeolites having a coordinated metal oxide to silica 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 alumina molar ratio of about 25:1 to 70:1, 
suitably modified. 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 
These siliceous zeolites may be employed in their acid forms, ion exchanged 
or impregnated with one or more suitable metals, such as Ga, Pd, Zn, Ni, 
Co and/or other metals of Periodic Groups III to VIII. The zeolite may 
include a hydrogenation-dehydrogenation component (sometimes referred to 
as a hydrogenation component) which is generally one or more metals of 
group IB, IIB, IIIB, VA, VIA or VIIIA of the Periodic Table (IU), 
especially aromatization metals, such as Ga, Pd, etc. Useful hydrogenation 
components include the noble metals of Group VIIIA, especially platinum, 
but other noble metals, such as palladium, gold, silver, rhenium or 
rhodium, may also be used. Base metal hydrogenation components may also be 
used, especially nickel, cobalt, molybdenum, tungsten, copper or zinc. The 
catalyst materials may include two or more catalytic components, such as a 
metallic oligomerization component (e.g., ionic Ni.sup.+2, and a 
shape-selective medium pore acidic oligomerization catalyst, such as ZSM-5 
zeolite) which components may be present in admixture or combined in a 
unitary bifunctional solid particle. 
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. 
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. In order to obtain the desired particle size for fluidization, 
the zeolite catalyst crystals are bound with a suitable inorganic oxide, 
such as silica, alumina, etc. to provide a zeolite concentration of about 
5 to 95 wt. %. In the description of preferred embodiments a 25% H-ZSM-5 
catalyst contained within a silica-alumina matrix and having a fresh alpha 
value of about 80 is preferred. 
The conversion of aliphatic oxygenates such as methanol and dimethylether 
to olefins is a process well-known in the art. This methanol-to-olefins 
("MTO") type process can be optimized to produce a major fraction of 
C.sub.2 -C.sub.4 olefins. The C.sub.3 + olefins are a valuable feed to the 
industrial process known as Mobil Olefins to Gasoline/Distillate ("MOGD"). 
Ethene is not as easily converted to more valuable hydrocarbons, and 
therefore it is advantageous to separate ethene from an olefinic feedstock 
for the MOGD process. 
It is a purpose of the present invention to upgrade ethene-rich streams to 
more easily processed C.sub.3 + olefinic materials. 
In its simplest form, the present process comprises contacting a C.sub.1 
-C.sub.4 aliphatic oxygenate feedstock under oxygenate conversion 
conditions in a primary fluidized bed reaction system with a crystalline 
medium pore shape selective zeolite catalyst; withdrawing a product 
effluent comprising ethene and C.sub.3 + olefins; separating the product 
effluent into a primary bottoms stream comprising C.sub.3 + olefins and a 
primary overhead stream comprising ethene; and contacting the primary 
overhead stream comprising ethene under high severity oligomerization 
conditions in a secondary fluidized bed reaction system with a crystalline 
medium pore shape selective zeolite catalyst. 
The process further comprises withdrawing from the secondary fluidized bed 
reaction system a product comprising unreacted ethene and C.sub.3 + 
olefins; and adding at least a portion of withdrawn product to the primary 
fluidized bed reaction system. 
In the secondary fluidized bed reaction severity conditions can be 
controlled to optimize yield of C.sub.4 -C.sub.9 aliphatic hydrocarbons. 
It is understood that aromatics and light paraffin production is promoted 
by those zeolite catalysts having a high concentration of Bronsted acid 
reaction sites. Accordingly, an important criterion is selecting and 
maintaining catalyst inventory to provide either fresh catalyst having 
acid activity or by controlling catalyst deactivation and regeneration 
rates to provide an apparent average steady state alpha value of about 2 
to 20. 
Reaction temperatures and contact time are also significant factors in the 
reaction severity, and the process parameters are followed to give a 
substantially steady state condition wherein the reaction severity index 
(R.I.) is maintained to yield a desired weight ratio of propane to 
propene. It is preferred to operate the steady state fluidized bed unit to 
hold the R.I. at about 0.1:1 to 1:1. While reaction severity is 
advantageously expressed as the weight ratio of propane:propene in the 
gaseous phase, it may also be approximated by the analogous ratios of 
butanes:butenes, pentanes:pentenes, or the average of total reactor 
effluent alkanes:alkenes in the C.sub.3 -C.sub.5 range. The optimum value 
will depend upon the exact catalyst composition, feedstock and reaction 
conditions. 
Particle size distribution can be a significant factor in achieving overall 
homogeneity in turbulent regime fluidization. It is desired to operate the 
process with particles that will mix well throughout the bed. Large 
particles having a particle size greater than 250 microns should be 
avoided, and 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. Accordingly, the fluidization regime is 
controlled to assure operation between the transition velocity and 
transport velocity. Fluidization conditions are substantially different 
from those found in non-turbulent dense beds or transport beds. Turbulent 
regime fluidization conditions for the conversion of ethylene are 
disclosed in U.S. Pat. No. 4,746,762 (Avidan et al), incorporated herein 
by reference. 
Several useful parameters contribute to fluidization in the turbulent 
regime in accordance with the process of the present invention. When 
employing a ZSM-5 type zeolite catalyst in fine powder form 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 apparaent "outside" volume) in the 
range from 0.6-2 g/cc, preferably 0.9-1.6 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 about 20 and 100 microns, preferably 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 fluidized bed where 
the superficial fluid velocity is 0.3-2 m/s, operation in the turbulent 
regime is obtained. Those skilled in the art will appreciate that at 
higher pressures, a lower gas velocity may be employed to ensure operation 
in the turbulent fluidization regime. 
Under optimized process conditions the turbulent bed has a superficial 
vapor velocity of about 0.3 to 2 meters per second (m/sec). At higher 
velocities entrainment of fine particles may become excessive and beyond 
about 3 m/sec the entire bed may be transported out of the reaction zone. 
At lower velocities, the formation of large bubbles or gas voids can be 
detrimental to conversion. Even fine particles cannot be maintained 
effectively in a turbulent bed below about 0.1 m/sec. 
A convenient measure of turbulent fluidization is the bed density A typical 
turbulent bed has an operating density of about 100 to 500 kg/m3, 
preferably about 300 to 500 kg/m.sup.3, measured at the bottom of the 
reaction zone, becoming less dense toward the top of the reaction zone, 
due to pressure drop and particle size differentiation. Pressure 
differential between two vertically spaced points in the reactor column 
can be measured to obtain the average bed density at such portion of the 
reaction zone. For instance, in a fluidized bed system employing ZSM-5 
particles having an apparent packed density of 750 kg/m.sup.3 and real 
density of 2430 kg/m.sup.3, an average fluidized bed density of about 300 
to 500 kg/m.sup.3 is satisfactory. 
By virtue of the turbulence experienced in the turbulent regime, gas-solid 
contact in the catalytic reactor is improved, providing at least 70% 
ethene conversion, enhanced selectivity and temperature uniformity. One 
main advantage of this technique is the inherent control of bubble size 
and characterisitic bubble lifetime. Bubbles of the gaseous reaction 
mixture are small, random and short-lived, thus resulting in good contact 
between the gaseous reactants and the solid catalyst particles. 
In an alternative embodiment, the process comprises contacting a C.sub.1 
-C.sub.4 aliphatic oxygenate feedstock in a primary fluidized bed reaction 
zone of a riser reaction system with a shape selective medium pore 
crystalline metallosilicate catalyst under oxygenate conversion conditions 
to obtain a primary effluent comprising ethylene and C.sub.3 + olefins; 
withdrawing the primary effluent; separating the primary effluent into a 
gaseous stream rich in ethylene and a liquid stream comprising C.sub.3 + 
olefins; contacting the liquid stream comprising C.sub.3 + olefins with a 
shape selective medium pore crystalline metallosilicate catalyst in a 
fixed bed reaction zone at elevated temperature and pressure to obtain a 
product stream comprising hydrocarbons boiling in the gasoline and 
distillate range; contacting the ethylene-rich stream with a shape 
selective medium pore crystalline metallosilicate catalyst in a secondary 
fluidized bed reaction zone of the riser reaction system under 
oligomerization conditions to obtain a secondary effluent comprising 
unreacted ethylene and heavier hydrocarbons; withdrawing the secondary 
effluent; and combining the secondary effluent with the C.sub.1 -C.sub.4 
aliphatic oxygenate feedstock for further conversion in the primary 
fluidized bed reaction zone. 
The above-described riser reaction system allows for a mixed feed to the 
oxygenates conversion reaction zone. The effluent from the oligomerization 
reaction zone comprises C.sub.3 + olefins, C.sub.6 + aromatics and 
unreacted ethene which function as reactive components in the aliphatic 
oxygenates conversion reactor. 
The present invention provides a process for integrating catalyst flow in a 
multi-stage system. The process comprises maintaining a primary fluidized 
bed reactor stage at oxygenate conversion conditions, the reactor 
containing solid catalyst particles comprising medium pore shape selective 
crystalline aluminosilicate; maintaining a secondary fluidized bed reactor 
stage at olefins oligomerization conditions, the reactor containing solid 
catalyst particles comprising medium pore shape selective crystalline 
aluminosilicate; withdrawing from the secondary fluidized bed reactor 
stage at least a portion of partially deactivated solid catalyst 
particles; and adding the partially deactivated solid catalyst particles 
to the primary fluidized bed reactor stage. 
Preferably, the solid catalyst particles in the secondary fluidized bed 
reactor stage have an average steady state alpha value of about 2 to 20 
and the solid catalyst particles in the primary fluidized bed reactor 
stage have an average steady state alpha value of about 1 to 10. 
In situations where the two fluidized bed reactors are operating at 
pressures which are relatively close to each other, the solid catalyst 
particles in both reactor stages are preferably passed to a common 
regenerator and catalyst stripper. Regenerated catalyst particles are then 
withdrawn from the common regenerator and passed to both the primary and 
secondary fluidized bed reactor stages. 
The feature of a common regenerator for catalyst particles in both reactor 
stages is conveniently incorporated into the present process. In its 
simplest form, the process comprises preparing a feedstock for a fixed bed 
olefins conversion to gasoline and distillate reactor comprising 
maintaining a reaction system comprising a lower section operatively 
connected to an upper section, the lower section comprising a secondary 
fluidized bed containing catalyst particles comprising shape selective 
medium pore crystalline metallosiliate catalyst and the upper section 
comprising a primary fluidized bed containing catalyst particles 
comprising shape selective medium pore crystalline metallosilicate 
catalyst; contacting a vapor stream comprising ethene with fluidized 
catalyst particles in the lower section to obtain an effluent comprising 
unreacted ethene and heavier hydrocarbons; withdrawing the effluent from 
the lower section; adding a stream comprising a C.sub.1 -C.sub.4 oxygenate 
or mixtures thereof to the uperr section; adding the lower section 
effluent with fluidized catalyst particles to the upper section to produce 
oligomerizable C.sub.3 + olefins. 
The lower section effluent containing fluidized catalyst particles can be 
added to the upper section of the reaction system either above the apex of 
the primary fluidized bed or directly into the primary fluidized bed. If 
further conversion is desired then the lower section effluent is contacted 
with catalyst in the primary fluidized bed. Otherwise the lower section 
effluent can be added above the apex of said fluidized bed so as to avoid 
further conversion. In both cases catalyst fines entrained in lower 
section effluent will be captured by cyclones in the oxygenates conversion 
reactor. 
The process further comprises withdrawing the feedstock comprising C.sub.3 
+ olefins from the riser reaction system; and adding the feedstock to a 
fixed bed reactor system maintained under conditions for the conversion of 
lower molecular weight C.sub.3 + olefins to hydrocarbons boiling in the 
gasoline and/or distillate range and containing catalyst particles 
comprising shape selective medium pore crystalline metallosilicate 
catalyst. From the fixed bed reactor system there is withdrawn a product 
comprising hydrocarbons boiling in the gasoline and/or distillate range. 
Catalyst inventory in the lower section of the reaction system is 
integrated with catalyst inventory in the upper section of the system by a 
process comprising withdrawing at least a portion of the catalyst 
particles in admixture with the effluent from the lower section; adding 
the withdrawn catalyst particles without separation from effluent to the 
upper section; withdrawing at least a portion of catalyst particles from 
the upper section; adding the particles withdrawn from the upper section 
to a catalyst regeneration system; regenerating the catalyst particles to 
obtain a reactivated catalyst; and returning reactivated catalyst to the 
lower section. The reactivated catalyst is lifted into the reaction system 
by the action of an ethene-rich feedstream which is forced upwardly 
through the lower section of the riser reaction system. A portion of the 
oxygenate feedstream can be combined with the ethene-rich feedstream to 
function as a lift gas for the reactivated catalyst. 
In a preferred embodiment, the crystalline metallosilicate catalyst 
particles in both the secondary fluidized bed of the lower section and the 
primary fluidized bed of the upper section comprise an aluminosilicate 
zeolite having the structure of ZSM-5. Most preferably, the zeolite 
comprises ZSM-5. 
The continuous multi-stage process comprises contacting the ethylene-rich 
stream in a high severity riser reactor at elevated temperature and 
pressure with a zeolite oligomerization catalyst having an average alpha 
value of at least about 2 to provide a hydrocarbon effluent comprising 
heavier hydrocarbons and unreacted light olefin; combining the hydrocarbon 
effluent with a stream comprising C.sub.1 -C.sub.4 aliphatic oxygenates to 
obtain a feedstock for an oxygenate conversion reactor; and contacting the 
combined feedstock in a fluidized bed oxygenate conversion reactor with a 
zeolite oxygenate conversion catalyst to obtain a product comprising 
C.sub.3 + oligomerizable olefins. The C.sub.3 + oligomerizable olefins are 
contacted with a medium pore shape selective crystalline aluminosilicate 
catalyst under conversion conditions to obtain a product comprising 
hydrocarbons boiling in the gasoline and/or distillate range, which are 
then withdrawn as product. 
In a preferred embodiment, the high severity riser reactor is operated at 
substantially the same pressure as the fluidized bed oxygenate conversion 
reactor. This allows for the use of a common regenerator, spent catalyst 
stripper and fines recovery section for the MOG and MTO processes. A 
typical MTO operation is conducted over a bed of HZSM-5 catalyst at about 
340 kPa (50 psia) at 515 degrees C. (960 degrees F.) at a space velocity 
(WHSV) of 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 2.26 
Ethylene 8.65 
Ethane .59 
Propylene 22.68 
Propane 4.09 
C.sub.4 hydrocarbons 
22.54 
C.sub.5 hydrocarbons 
12.42 
C.sub.6 hydrocarbons 
7.36 
C.sub.7 hydrocarbons 
5.99 
C.sub.8 + hydrocarbons 
13.41 
Unknown HC as C.sub.5 
.01 
______________________________________ 
Applicants have discovered a continuous multistage fluidized bed catalytic 
process for converting a C.sub.1 -C.sub.4 aliphatic oxygenate feedstock to 
liquid hydrocarbon products comprising: maintaining a primary fluidized 
bed of finely divided acid medium pore zeolite catalyst particles in 
steady state equilibrium fluidized condition under oxygenate conversion 
reaction conditions, wherein the average catalyst particle acid value is 
maintained at a first predetermined value to optimize conversion of 
oxygenate feedstock to lower olefins comprising ethylene and C.sub.3 + 
olefins; contacting a C.sub.1 -C.sub.4 aliphatic oxygenate feedstock with 
acid zeolite catalyst particles in the primary fluidized bed to produce a 
lower olefin effluent; withdrawing the lower olefin effluent and 
separating the effluent to obtain a first effluent fraction rich in 
ethylene and a second effluent fraction rich in C.sub.3 + olefins; 
maintaining a secondary fluidized bed of finely divided acid medium pore 
zeolite catalyst particles in steady state equilibrium fluidized condition 
under reaction severity conditions sufficient to convert a major amount of 
ethylene, wherein the average catalyst particle acid value is maintained 
at a second predetermined acid value higher than said first value by 
addition of fresh acid medium pore zeolite catalyst particles to optimize 
conversion of ethylene to C.sub.5 + hydrocarbons; contacting the first 
effluent fraction rich in ethylene with acid zeolite catalyst particles in 
the secondary fluidized bed to obtain a C.sub.5 +hydrocarbon effluent; 
withdrawing acid catalyst particles from the secondary fluidized bed; 
adding a first portion of withdrawn secondary fluidized bed particles to 
the primary fluidized bed; adding a second portion of withdrawn secondary 
fluidized bed particles to a catalyst regeneration zone; withdrawing 
deactivated acid catalyst particles from the primary fluidized bed to 
maintain equilibrium conditions therein; adding withdrawn primary 
fluidized bed particles to the catalyst regeneration zone; oxidatively 
regenerating catalyst particles from said primary and secondary fluidized 
beds in the catalyst regeneration zone; withdrawing regenerated catalyst 
from the oxidative regeneration zone; and adding regenerated catalyst to 
at least the secondary fluidized bed. 
An apparatus useful for increasing the yield of liquid hydrocarbons from a 
C.sub.1 -C.sub.4 aliphatic oxygenate feedstock is disclosed. The apparatus 
comprises a primary fluidized bed reaction zone containing shape selective 
medium pore crystalline metallosilicate catalyst; a primary separation 
zone for separating a primary effluent into a gaseous stream rich in 
ethylene and a liquid stream comprising C.sub.3 + olefins; a means for 
passing said primary effluent from primary reaction zone to primary 
separation zone; a fixed bed reaction zone containing shape selective 
medium pore crystalline metallosilicate catalyst; a means for passing the 
liquid stream from the primary separation zone to the fixed bed reaction 
zone; a secondary fluidized bed reaction zone containing shape selective 
medium pore crystalline metallosilicate catalyst; a means for passing the 
gaseous stream from the primary separation zone to the secondary reaction 
zone; a secondary separation zone for separating a secondary effluent into 
a gaseous stream comprising ethylene and a liquid stream comprising 
heavier hydrocarbons; a means for passing said secondary effluent from 
secondary reaction zone to secondary separation zone; a means for passing 
at least a portion of liquid stream from the secondary separation zone to 
the fixed bed reaction zone; a means for withdrawing from the fixed bed 
reaction zone a product stream comprising hydrocarbons boiling in the 
gasoline and distillate range; a common catalyst regeneration zone for 
oxidative regeneration of deactivated catalyst from both primary and 
secondary reaction zone; a means for passing deactivated catalyst from 
primary reaction zone to catalyst regeneration zone; a means for passing a 
first portion of deactivated catalyst from secondary reaction zone to said 
primary reaction zone; a means for passing a second portion of deactivated 
catalyst from secondary reaction zone to said catalyst regeneration zone; 
and a means for passing reactivated catalyst from common catalyst 
regeneration zone to both the primary and secondary fluidized bed reaction 
zones. 
An apparatus is disclosed for converting a C.sub.1 -C.sub.4 aliphatic 
oxygenate feedstock to hydrocarbon products. The apparatus comprises a 
lower section comprising a secondary fluidized bed containing shape 
selective crystalline zeolite catalyst particles having an apparent 
average steady state alpha value of about 2 to 20; an operatively 
connected upper section comprising a primary fluidized bed containing 
shape selective crystalline zeolite catalyst particles having an apparent 
average steady state alpha value of about 1 to 10; a means for adding an 
ethene-containing feedstream to the lower section; a means for passing a 
secondary effluent containing C.sub.3 + olefins, unreacted ethene and 
catalyst particles from the lower section to the upper section, said means 
either adding the secondary effluent directly to the primary fluidized bed 
for contact with zeolite catalyst or adding the secondary effluent above 
the uppermost level of the primary fluidized bed for admixture with a 
primary effluent; a means for adding a feedstream comprising a C.sub.1 to 
C.sub.4 oxygenate or mixtures thereof to the primary fluidized bed of the 
upper section; a means for withdrawing a primary effluent from the upper 
section; a common catalyst regeneration zone for oxidative regeneration of 
deactivated catalyst from both the primary and secondary fluidized beds; a 
means for passing deactivated catalyst from upper level to catalyst 
regeneration zone; a means for passing reactivated catalyst from the 
regeneration zone to the ethene-containing feedstream means; and a means 
for adding fresh catalyst to the ethene-containing feedstream means. 
Alternatively, the apparatus further comprises a means for passing a 
portion of said oxygenate feedstream to the ethene-containing feedstream 
means. 
Referring to FIG. 1, a C.sub.1 -C.sub.4 oxygenate feedstock, preferably a 
mixture of methanol and dimethyl ether, is passed as by line 1 to a 
primary fluidized bed oxygenates conversion reactor 2 containing a shape 
selective medium pore crystalline metallosilicate catalyst. The reactor is 
maintained at a conversion temperature of about 450 degrees C. to 532 
degrees C. and a pressure of about 69 KPa to 900 KPa. The amount of 
oxygenate converted per pass is about 100%. 
The crystalline metallosilicate catalyst has an average alpha value of 
about 1 to 10. The catalyst preferably comprises an aluminosilicate 
zeolite having the structure of ZSM-5. In a most preferred embodiment, the 
catalyst comprises ZSM-5. 
An effluent comprising unreacted oxygenate, ethene, C.sub.3 + olefins, 
ethane, and other light gases is withdrawn from the oxygenates conversion 
reactor as by line 3. The entire effluent is conducted via line 3 to a 
separation unit 4 wherein a vapor stream containing a major amount of 
ethene exits overhead via line 7, and a liquid hydrocarbon stream 
comprising C.sub.3 + olefins such as propene, n-butene, isobutene, the 
isomeric pentenes, and hexene isomers is withdrawn via line 5. 
The ethene-rich stream, which contains a significant quantity of 
unrecoverable C.sub.3 + olefins which are also upgraded to gasoline range 
hydrocarbons, enters a secondary fluidized bed reaction zone 6 maintained 
under olefins oligomerization conditions to upgrade the ethene to a 
C.sub.5 -C.sub.10 hydrocarbon product boiling in the gasoline range. The 
fluidized bed contains catalyst comprising shape selective medium pore 
crystalline metallosilicate particles. In a preferred embodiment the 
catalyst particles are aluminosilicate materials such as zeolites. Most 
preferably the catalyst is a zeolite having the structure of ZSM-5 with an 
average alpha value of about 2 to 50. 
The fluidized bed reaction zone 6 is maintained at an olefins 
oligomerization temperature of about 232 degrees C. to 532 degrees C. and 
a pressure of about 69 KPa to 2,061 KPa. The amount of ethene converted 
per pass is more than about 70%, preferably about 93 to 96%. 
A reaction effluent comprising unreacted ethene, C.sub.5 -C.sub.10 
hydrocarbons boiling in the gasoline range, and light gases is withdrawn 
from fluidized bed reaction zone 6 and passed as by line 11 to a 
separation column 10. An overhead containing unreacted ethene is withdrawn 
from the separator and sent to fuel main via line 9. A bottoms stream 
comprising C.sub.5 -C.sub.10 hydrocarbons boiling in the gasoline range is 
withdrawn from the separation column 10 and passed via line 12 to a fixed 
bed reaction zone 8. The bottoms stream is optionally passed directly to 
the gasoline pool. 
A second feed to fixed bed reaction zone 8 is the bottoms stream obtained 
from separation unit 4. The bottoms stream comprising C.sub.3 + olefins 
enters the fixed bed reactor 8 as by line 5. 
The fixed bed reaction zone 8 contains a shape selective medium pore 
crystalline metallosilicate catalyst which has an average alpha value of 
about 3 to 80. The catalyst preferably comprises an aluminosilicate 
zeolite having the structure of ZSM-5. In a most preferred embodiment, the 
catalyst comprises ZSM-5. 
The fixed bed reactor is maintained at a temperature of about 204 degrees 
C. to 315 degrees C. and a pressure of about 2,747 KPa to 20,604 KPa. A 
higher pressure is employed to obtain lube components in the reaction 
product. 
When C.sub.3 + olefins contact the metallosilicate catalyst contained in 
the fixed bed reactor, they are oligomerized to produce hydrocarbons 
boiling in the gasoline and/or distillate range. A crude product stream 
comprising gasoline and/or distillate hydrocarbons is withdrawn from the 
fixed bed reactor via line 13. The product stream can be purified in a 
series of separation steps or sent to a reaction zone for further 
processing. In a preferred embodiment, ethene-containing light components 
from the fixed bed reaction zone are separated in fractionation column 14 
and returned as by line 17 to zone 6 in order to maximize C.sub.5 + 
hydrocarbon yield. The light components optionally contain C.sub.3 
-C.sub.4 hydrocarbons. Hydrocarbon product containing C.sub.5 + 
hydrocarbons, including gasoline and optionally distillate and lube 
components, is withdrawn as by line 18. 
In an alternative embodiment, the reaction effluent withdrawn from 
fluidized bed reaction zone 6 via line 11 is passed as by line 15 to the 
fluidized bed oxygenates conversion zone 2. The effluent from reaction 
zone 6 can be admixed with oxygenates feedstock prior to entering 
fluidized bed 2. Alternatively, at least a portion of reaction effluent 
withdrawn from secondary zone 6 as by line 11 can be returned to 
separation unit 4 via line 16. 
With regard to catalyst flow, partially deactivated catalyst particles are 
withdrawn from secondary fluidized bed 6 as by line 21. A first portion of 
said catalyst particles is added to the primary fluidized bed via line 22 
to preserve equilibrium conditions in the oxygenates conversion reactor. A 
second portion of said withdrawn catalyst particles is optionally added to 
catalyst regeneration unit 20 via line 23 where catalyst is regenerated as 
by oxidative means. The regeneration unit 20 also serves to reactivate 
catalyst particles which are withdrawn from primary fluidized bed 2 via 
line 24. 
From the common catalyst regenerator 20 is withdrawn reactivated catalyst 
particles which are added to either one or both fluidized beds. Catalyst 
particles can be added to the primary fluidized bed as by line 25 and to 
the secondary fluidized bed via line 26. In a preferred embodiment, 
regenerated catalyst is added to the secondary fluidized bed 6 along with 
catalyst make-up, which is added as by line 28. Catalyst make-up can be 
added to primary fluidized bed 2 via line 27. Catalyst purge is withdrawn 
from regenerator 20 via line 29. 
FIG. 2 shows a reacton system comprising a fluidized bed olefins upgrading 
riser reactor 6 and an operatively connected fluidized bed oxygenates 
conversion reactor 2. The reaction system permits successful upgrading of 
an ethene-rich stream which is a byproduct from conversion of C.sub.1 
-C.sub.4 oxygenates to olefins. In another embodiment an ethene-rich 
stream obtained from any refinery source is employed as a feed to the 
riser reactor. 
Referring to FIG. 2, a C.sub.1 -C.sub.4 oxygenate feedstock, preferably an 
equilibrium mixture of methanol and dimethylether, is passed via line 1 to 
fluidized bed in reaction zone 2 where it contacts finely divided solid 
catalyst particles comprising a shape selective medium pore crystalline 
metallosilicate. The metallosilicate preferably comprises an 
aluminosilicate zeolite having the structure of ZSM-5. Most preferably, 
the catalyst comprises ZSM-5. 
A product comprising C.sub.3 + oligomerizable olefins is withdrawn from 
reactor 2 and passed to a separation unit (not shown). From separation 
unit is withdrawn an overhead stream comprising ethene and a bottoms 
stream containing C.sub.3 + olefins. 
The bottoms stream rich in C.sub.3 + olefins is an excellent feedstock for 
a catalytic reaction system maintained under conditions for oligomerizing 
olefins to gasoline, distillate and lube range hydrocarbons. The C.sub.3 + 
olefins-rich bottoms stream can be passed directly to the catalytic 
reaction system or mixed with other olefinic feedstock prior to entering 
the system. 
The reaction system for producing gasoline, distillate and lube range 
hydrocarbons contains a shape selective medium pore crystalline 
metallosilicate catalyst having an average steady state alpha value of 
about 3 to 80. The catalyst preferably comprises an aluminosilicate 
zeolite having the structure of ZSM-5. In a most preferred embodiment, the 
catalyst comprises ZSM-5. From the reaction system is withdrawn a product 
comprising hydrocarbons boiling in the gasoline and/or distillate range. 
An ethene-containing overhead stream from the separation unit (not shown) 
enters secondary stage fluidized bed riser reactor 6 as by line 3. In an 
alternative embodiment, a portion of oxygenate feed is added to the 
ethene-containing stream as by line 16. Ethene contacts solid catalyst 
particles comprising a crystalline shape selective medium pore 
metallosilicate, preferably ZSM-5 zeolite, under high severity reaction 
conditions. The catalyst has an average steady state alpha value of about 
2 to 20. At least a portion of the ethene is catalytically converted to 
C.sub.5 -C.sub.10 hydrocarbons boiling in the gasoline range. 
An effluent comprising C.sub.5 -C.sub.10 hydrocarbons and unreacted ethene 
is withdrawn from riser reactor 6 as by line 13 and admixed with the 
reactor effluent via line 18 or the reactor bed via line 14 in primary 
stage reaction zone 2. Oxygenate enters reactor 2 via line 1. The apex or 
uppermost level of fluidized bed in reactor 2 is shown by segmented line 
15. 
The light hydrocarbon stream recovered from the oxygenates conversion stage 
via line 5 preferably contains a major amount of C.sub.2 -C.sub.4 olefins. 
The novel system includes a separation unit for recovering ethene from the 
primary stage olefinic vapor effluent including a sorption tower 
operatively connected to selectively sorb C.sub.3 + hydrocarbons from the 
olefinic vapor effluent in a liquid sorption stream. Since the separation 
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 + sorbate are 
provided. 
A suitable sorption fractionation system is described in U.S. patent 
application Ser. No. 508,779, filed June 29, 1983 (Hsia et al), now U.S. 
Pat. No. 4,479,812 the disclosure of which is incorporated herein by 
reference. The C.sub.2 - and C.sub.3 + 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 - 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, C.sub.2 H.sub.6 and CH.sub.4) from the C.sub.3 + hydrocarbon. 
The gasoline sorbent is an aliphatic hydrocarbon mixture boiling in the 
normal gasoline range of about 50 to 165 C. (125 to 330 F.), with minor 
amounts of C.sub.4 -C.sub.5 alkanes and alkenes. 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 + hydrocarbons in the feedstock, with optimum 
operation utilizing a sorbert:sorbate molar ratio about 1:1 to 1.5:1. 
In the process design of FIG. 2 catalyst flow is upward through riser 
reactor 6 and is passed to the turbulent regime fluidized bed reactor 2. 
Partially deactivated catalyst, including both oligomerization catalyst 
and oxygenates conversion catalyst, is withdrawn as by line 7, lifted by 
air as by line 8, and contacted with an oxygen-containing gas in a common 
catalyst regeneration zone 4 for oxidative regeneration. Regenerated 
catalyst is returned to riser reactor 6 via line 10, and catalyst make-up 
is added to reactor 6 via line 11. The daily catalyst make-up is about 0.3 
wt. % of the total catalyst inventory. A catalyst purge is withdrawn from 
regeneration unit 4 via line 9. Flue gas is withdrawn from regenerator 4 
as by line 12. 
An additional feature of the present design as shown in FIG. 2 is a 
combined catalyst fines filtration system and process wherein the catalyst 
fines from both the primary turbulent regime fluidized bed oxygenates 
conversion reactor and the secondary olefins oligomerization riser reactor 
are filtered and recovered in a common system. The system comprises 
cyclones and/or a sintered metal filter system to recover fines from the 
combined reactors effluent stream. 
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 process for catalytic conversion of olefins to heavier hydrocarbons 
by catalytic oligomerization using an acid crystalline zeolite, such as a 
catalyst having the structure of ZSM-5, 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 80 to 99% in the distillate mode, only about 10 to 50% of the 
ethylene component will be consumed. Also, high-pressure pressure 
conversion of ethene is unfavorable because of the prohibitively high cost 
of compressing ethene-rich feedstocks. Accordingly, the ethene is 
advantageously recovered prior to the oligomerization stage. In a 
preferred embodiment the olefinic feedstock is obtained form lower 
aliphatic oxygenates, such as methanol or dimethylether (DME). A typical 
crude methanol feedstock may contain 4 to 17% water, with minor amounts of 
carbon oxides, methane, DME, etc. 
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. The oxygenate is fed 
continuously under low pressure where it is raised to process temperature, 
and introduced to the MTO reactor system. The initial dehydration reactor 
is followed by a catalytic reactor containing zeolite conversion catalyst. 
The effluent is cooled to condense water and a major amount of C.sub.5 - 
liquid hydrocarbons. These liquids are separated from the hydrocarbon 
vapor in a phase separator. Byproduct water may be recovered from 
unreacted feedstock and discarded. The liquid hydrocarbon phase and the 
ethene-rich light hydrocarbon vapor streams are recovered from a 
separator. A suitable fluid catalyst apparatus is disclosed in U.S. Pat. 
No. 4,379,123 (Daviduk and Haddad). 
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.