Note: Descriptions are shown in the official language in which they were submitted.
917~i9
UPC,RADI~, DIENE-CONTAINI~æ HYD~)CARBONS
This invention relates to a catalytic technique for
upgrading olefin streams rich in dienes to heavier hydrocarbons rich
in aliphatics and aromatics.
Developments in zeolite catalysis and hydrocarbon
conversion processes have created interest in utilizing olefinic
feedstocks for produci~g C5 gasoline, diesel fuel, etc. In
addition to basic chemical reactions promoted by ZSM-5 type zeolite
catalysts, a numbee of discoveries have contributed to the
development of new industrial processes. ~hese are safe,
environmentally acceptable processes for utilizing feedstocks that
contain olefins . Conversion of C2-C4 alkenes and alkanes to
produce aromatics-rich liquid hydrocarbon products were found by
Cattanach (US 3,760,024) and Yan et al (US 3,845,150) to be
effective processes using the ZSM -5 type zeolite catalysts. In
U.S. Patents 3,960,978 and 4,021~502, Plank, Rosinski and Givens
disclose conversion of C2-C5 olefins, alone or in admixture with
paraffinic components, into higher hydrocarbons over crystalline
zeolites having controlled acidity. Garwood et al. have also
contributed to the understanding of catalytic olefin upqrading
techniques and improved processes as in U.S. Patents 4,150,062,
4,211,640 and 4,227,992.
Conversion of 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 C5 aliphatic and aromatic hydrocarbons and C4
hydrocarbons, in particular iso-butane. Product distribution for
liquid hydrocarbons can be varied by controlling process conditions,
such as temperature, pressure and space velocity. Gasoline
~C5-C10) is readily formed at elevated temperature e.g., up to
about 700C. and moderate pressure from ambient to about 5500 kPa,
preferably about 200 to 2900 kPa. Olefinic gasoline can be produced
X.
~91~69
417~+ -2-
in good yield and may be recovered as a product or fed to a low
severity, high pressure reactor system for further conversion to
heavier distillate-range products. Distillate mode operation can be
employed to maximize production of Cl0 aliphatics by reacting
the lower and intermediate olefins at high pressure and moderate
temperature. Operating details for such oligomerization units are
disclosed in U.S. Patents 4,456,779; 4,497,968 (~en et al.) and
4,433,185 (Tabak). At moderate temperature and relatively high
pressure, the conversion conditions favor distillate-range product
having a normal boiling point of at least 165C. (330F.).
Many feedstocks of commercial interest, such as thermal
cracking byproduct, etc., contain both mono-olefins and diolefins
(e.g. C2-C6 mono-alkenes and C4+ dienes) along with
Cl-Cl0 light aliphatics, and a minor amount of aromatics.
Gaseous and liquid streams containing dienes are typically produced
in thermal cracking operations. One common example is pyrolysis
gasoline, which is produced as ethene (ethylene) byproduct. Such
diene-containing streams are often difficult to process due to poor
thermal stability and the tendency of dienes to form coke and qum
deposits. m is complicates preheating of such streams into the high
temperatures required of most catalytic upgrading processes. Prior
attempts to upgrade such materials have pretreated the feedstock to
hydrogenate the diolefin selectively, as in U.S. Patent No.
4,052,477 (Ireland et al). m e present invention is concerned with
providing a safe and low cost technique for catalytically converting
diene-rich streams to high value C4+ products
It has been found that diene-containing olefinic light
hydrocarbons can be upgraded directly to liquid hydrocarbons rich in
C5+ aliphatics and aromatics by catalytic _onversion. This
technique is particularly useful for upgrading C4+ liquid
pyrolysis products, which may contain minor amounts of ethene,
propene, C2-C4 paraffins and hydrogen produced in cracking
petroleum fractions, such as naphtha, ethane or the like. By
upgrading the complex olefinic by-product, gasoline yield and/or
~,91~69
417~+ -3-
aromatics production of cracking units can be significantly
increased.
Accordingly, the present invention provides a process for
upgrading diene-rich liquid olefinic feed comprising at least 1 wt
diene and a total C4-C6 olefin content of ahout 5 to 90 wt % to
liquid product with a reduced diene content characterized by
maintaining a relatively dense phase fluidized bed of zeolite
catalyst at 220 to 510C; adding the feed at a temperature lower
than the bed temperature into the fluidized bed; and converting a
majority of dienes in the feed; and recovering hydrocarbon product
containing a major amount of C4+ hydrocarbons and having a C3
to C5 alkane:alkene weight ratio of 0.2:1 to 200:1.
FIG. 1 is a schematic view of a fluidized bed reactor
system of the present invention;
FIG. 2 is a vertical cross section view of a liquid-gas
feed nozzle which is employed to introduce low temperature diene
feed into the feactor bed;
FIG. 3 is an aging plot showing the effect of adding
butadiene to a C2-C4 olefinic feed.
FIG. 4 is a schematic view of a fluidized bed reactor and
regeneration system of the invention;
FIG. 5 is a process flow sheet for converting olefin
feedstock to aromatics-rich product, showing a reactor and effluent
separation equipment.
Catalysts
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 A1, Ga, B, Fe or mixtures thereof, within the zeolitic
framework. m ese medium pore zeolites are favored for acid
~;?..91~
~178+ -4-
catalysis; ho~ever, the advantaqes of %SM-5 structures may be
utilized by employing highly siliceous materials or cystalline
metallosilicate having one or more tetrahedral species havinq
varying degrees of acidity. ZSM-5 crystalline structure is readily
recognized by its X-ray diffraction pattern, which is described in
U.S. Patent No. 3,702,866 (Argauer, et al.).
m e oligomerization catalysts preferred for use herein
include the medium pore (i.e., about 5-7A) zeolites having a
silica-to-alumina ratio of at least 12, a constraint index of 1 to
12 and acid cracking activity (alpha value) of 10-250, more
preferably 10 to ~0 based on total catalyst weight. In the
fluidized bed reactor the coked catalyst may have an apparent
activity (alpha value) of 10 to 80 under the process conditions to
achieve the required degree of reaction severity.
Representative zeolites are ZSM-5, ZSM-ll, ZSM-12, ZSM-22,
ZSM-23, ZSM-35, ZSM-38, and ZSM-48. Other suitable zeolites are
disclosed in U.S. Patents 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. 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 use ZSM-5 having a
silica:alumina molar ratio of about 25:1 to 70:1, suitably modified
if desired to adjust acidity and aromatization characteristics.
mese 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-dehydroqenation
component (sometimes eeferred 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 (IUPAC), 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
~ ?.9~7fi9
41~S+ 5
rhodium, may also be used. Rase metal hydrogenation components may
also be used, es~eeially 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+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. It is possible to utilize an ethene dimerization
metal or oligomerization agent to effectively convert feedstock
ethene in a continuous reaction zone.
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 O.Ol to 2 microns or more.
In order to obtain the desired particle size for fluidization in the
turbulent regime, the zeolite catalyst crystals are preferably bound
with a suitable inorganic oxide, such as silica, alumina, etc. to
provide a zeolite concentration of about 5 to 95 wt. %. It is
advantageous to employ ZSM-5 having a silica:alumina molar ratio of
25:1 or greater in a once-through fluidized bed unit to convert 60
to lO0 percent, preferably at least 75 wt ~, of the monoalkenes and
dienes in the feedstock. In the description, a 25~ H-ZSM-5 catalyst
calcined with 75% silica-alumina matrix binder is employed unless
otherwise stated.
~IIlIDIZED BED REP~CTION ZONE
30 It is essential that the diolefin containing feed contact a
fluidized bed of catalyst. Fixed catalyst beds will not work in the
process of the present invention.
~9~769
4178+ -6-
Fluidization is a complex phenomenon. In general terms,
suitable fluidized beds include both dense phase beds and dilute
phase beds. The lower and uppee limits of fluidization will be
discussed.
In an expanded bed, there is enough of an increase in
up-flowing gas rate to cause particles to move apart and a few even
vibrate and move about in restricted regions. An expanded bed is
usually not suitable for use herein.
An ebullating bed represents the lower threshold of
fluidization required for use in the present invention. An
ebullating bed permits just enough circulation to work. Any gum
deposited on the catalyst might deactivate the catalyst near the
feed inlet momentarily, but there would be enough gross circulation
within the reactor so that incoming feed would see active catalyst.
m e gummed up catalyst could circulate to other portions of the
ebullating bed and, in time, the gum components would be cracked and
removed, in situ. A portion of the catalyst could also be
continuously withdrawn and replaced with fresh or regenerated
catalyst.
A conventional dense phase fluidized bed is preferred.
Gas-solid systems in this stage are also called aggregative
fluidized beds, a heterogeneously fluidized bed, a bubbling
fluidized bed, or simply a gas fluidized bed.
An especially preferred gas-fluidized bed is the turbulent
fluidized bed disclosed in US 4,547,616. Such a turbulent fluidized
bed is characterized by vigorous solid mixing, and strong
interactions between the gas and solid phases.
Dilute phase fluidiæed beds represent an upper limit on
fluidization and are also suitable for use herein. This type of 30 fluidized bed is also referred to as a disperse-, dilute-, or
lean-phase fluidized bed. Such beds oocur when the superficial
vapor velocity in the bed exceeds the terminal velocity of the
solids in the bed.
769
417~+ -7-
It is also possible, and may be very beneficial, to combine
one or more type of fluidized beds for use in the present
invelntion~ Specifically, all or a portion of the feed can be added
to a dilute phase riser which discharges into a dense phase bed. In
another embodiment liquid feed can be added to a relatively dense
phase bed over which is a dilute phase transpoet riser whch
discharges into another dense bed of catalyst or into
catalyst/product separation means.
In all embodiments, the feed addition means has a low
enough temperature/residence time so that plugging of the feed
distributor or feed nozzle is avoided, and feed sees an overwhelmin~
amount of hot, active catalyst. Undoubtedly, some dienes form gum
on the catalyst, but because the bed is fluidized no plugging of the
bed can occur. The very active zeolite catalyst will eventually
crack much of the gum, or gum precursor material, to lighter
products in the fluidized bed reactor.
Particle size distribution can be a significant factor in
fluidization. It is desired to operate the process with particles
that will mix well throughout the bed. Larqe 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 l to 150 microns. Average particle size is usually 20 to lO0
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 lO to 25 wt ~ of the total catalyst in the
reaction zone in the size range less than 32 microns. This size
range of fluidizable particles is classified as Geldart Group A.
m e fluidization regime is preferably controlled to assure operation
between the transition velocity and transport velocity, and these
fluidization conditions are substantially different from those found
in non-turbulent dense beds or transport beds.
~91~9
417~+ -8-
In the discussion that follows, percents are by weiqht
unless otherwise stated.
Feedstocks - Reaction Conditions
Suitable olefinic feedstocks comprise C4-C6 alkenes
including conjugated dienes such as 1,3-butadiene, pentadiene
isomees, hexadienes, cyclic dienes, or similar C4 aliphatic
liquid hydrocarbons having diethylenic conjugated unsaturation.
Aromatics coproduced with the liquid olefinic components may be
cofed or sepaeated by solvent extraction prior to conversion of the
diene-rich feedstock. Non-deleterious components, such as paraffins
and inert gases, may be present. A particularly useful feedstock is
a liquid by-product of pyrolysis or thermal cracking units
containing typically 40-95 wt ~ C4-C6 total mono-olefins and
di-olefins, including about 5-60 wt. % diene, along with varying
amounts of C3-C8 paraffins, aromatics and inerts. Specific
examples are given in Table 1 below. The process tolerates a wide
range of lower alkanes, from 0 to 95%. Preferred pyrolysis
feedstocks contain more than 50 wt % C4-C6 lower aliphatic
hydrocarbons, and contain sufficient olefins to provide an olefinic
partial pressure of at least 50 kPa. ~nder the high severity
reaction conditions employed in the present invention, lower alkanes
may be partially converted to heavier hydrocarbonds
The reaction severity conditions can be controlled to
optimize yield of C4-Cg h~drocarbons or of C6-C8 aromatics.
It is understood that aromatics and light paraffin production is
promoted by those æeolite catalysts having a high concentration of
Bronsted acid reaction sites. Accordingly, an important criterion
` is selecting and maintaining catalyst inventory to provide either
fresh or regenerated catalyst having the desired properties.
Typically, acid cracking activity (alpha value) can be maintained
from high activity values greater than 100 or 200 to significantly
lower values under steady state operation by controlling catalyst
X
~J9i7fi9
4178~ ~9~
deactivation and regeneration rates to provide an apparent average
alpha value below 100, preferably 10 to 80.
Reaction temperatures and contact time are also significant
factors in the reaction severity. The reaction severity index
(R.I.) is the weight ratio of propane to propene in reactor
effluent. While this index may vary from about 0.2 to 200, it is
preferred to operate the steady state fluidized bed unit to hold the
R.I. below about 50, with optimum operation at 0.7 to 2 in the
substantial absence of added propane. While reaction severity is
advantageously determined by 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 C3-C5 ranqe.
Accordingly, these alternative expressions may be a more accurate
measure of reaction severity conditions when propane is added to the
feedstock. m e optimal value will depend upon the exact catalyst
composition, feedstock and reaction conditions; however, the typical
diene-rich feed mixtures used in the examples herein and additional
olefinic feeds can be optimally upgraded to the desired
aliphatic-rich gasoline by keeping the R.I. at about 1.
Upgrading of olefins with hydrogen contributors in
fluidized bed cracking and oligomerization units is taught by Owen
et al in U. S. Patent 4,090,949. This technique is particulatly
useful for operation with a pyrolysis cracking unit to increase
overall production of liquid product.
m e use of fluidized bed catalysis permits the conversion
system to be operated at low pressure drop, which in an economically
practical operation can provide a maximum operating pressure only 50
to 200 kPa above atmospheric pressuee. Somewhat higher pressures,
up to 2500 kPa may be used to favor aromatics production. Close
temperature control is possible by turbulent regime operation. The
temperature can be maintained within close tolerances, often less
than 5C. Except for a small zone adjacent the bottom gas inlet,
~ ~917~i9
417S~ -10-
the midpoint measurement is representative of the entire bed, due to
the thorough mixing achieved.
Refereing now to FIG. 1, a reactor vessel 2 is shown
provided with heat exchange tube means 4. There may be several
separate heat exchange steam generating tube bundles so that
temperature control can be separately exercised over the fluid
catalyst bed. The bottoms of the tubes are spaced above a feed
distributor grid 8 sufficiently to be free of jet action by the
charged gas passinq through the small diameter holes in the grid 8.
Although depicted without baffles, the vertical reaction zone can
contain open end tubes above the grid for maintaining hydraulic
constraints, as disclosed in US Pat. 4,251,484 (Daviduk and
Haddad). Optionally, a variety of horizontal baffles may be added
to limit axial mixing in the reactor. Heat released from the
reaction can be controlled by adjusting feed temperature in a known
manner. A large portion of reaction heat can be removed by feedinq
cold liquid into the reactor at a temperature below averaqe bed
temperature preferably at least 200C below. In the reactor
configuration shown the heat exchanger tubes can function to limit
mixing in the reactor.
The system provides for withdrawing catalyst from above
grid 8 by conduit means 10. This flow line can be provided with
control valve means 12 for passage to catalyst regeneration invessel
13, where coked catalyst particles are oxidatively regenerated in
contact with air or other regeneration gas at high temperature. The
oxidatively regenerated catalyst is then passed to the reactor fluid
bed of catalyst by conduit means 14 and flow control valve lfi. The
eegenerated catalyst is charged to the catalyst bed sufficiently
below the upper interface to achieve good mixing in the fluid bed.
Since the flow of regenerated catalyst passed to the reactor can be
relatively small, hot regenerated catalyst does not ordinarily upset
the temperature constraints of the reactor operations in a
significant amount.
~?~ fi9
417~ -11-
Initial fluidization is achieved by forcinq a lift gas
upwardly through the catalyst, a light aliphatic C4 gas, with
or without diluent or recycle, may be charged through inlet port 20A
at a bottom portion of the reactor in open communication with
chamber 24 beneath grid 8. Pressurized feedstock is introduced
above reactant distributor grid 8 via supply conduit 21, pump 22 and
distributor conduit 23 to one or more spray nozzle means, described
and depicted in Fig. 2. The liquid is dispersed into the bed of
catalyst thereabove at a velocity sufficient to form a generally
upwardly flowing suspension of atomized liquid reactant with the
catalyst particles and lift gas.
Advantageously, the liquid diene-containing reactant feed
is injected into the catalyst bed by atomizing the pressurized
liquid feedstream to form readily dispersible liquid particles
having an average size of 300 microns or less. This contributes to
rapid vaporization of the liquid at process pressure. ~xothermic
conversion provides sufficient heat to vaporize the liquid quickly,
thus avoiding deleterious liquid phase reactions of the diene
components, which tend to form carbonaceous deposits such as heavy
coke, gums, etc.
A plurality of sequentially connected cyclone separator
means 30, 32 and 34 provided with diplegs 36, 38 and 40 respectively
are positioned in an upper portion of the reactor vessel comprisinq
dispersed catalyst phase 28.
The product effluent separated from catalyst particles in
the cyclone separating system then passes to a plenum chamber 4~
before withdrawal via conduit 46, operatively connect with effluent
separation system 50. The product effluent is cooled and separated
to recover C5+ liquid hydrocarbons, gaseous recycle or offgas,
along with any byproduct water or catalyst fines carried over.
portion of the light gas effluent fraction may be recycled by
compressing to form a motive gas for the liquid feed or via recycle
conduit 20B for use as lift gas. The recovered hydrocarbon product
comprising C5 olefins and/or ~ aromatics, paraffins and
91769
4178+
naphthenes is thereafter processed as required to provide a desired
gasoline or highe~ boiling product.
Referring now to Fig. 4, in which maximum BTX production is
sought, liquid feedstock is passed at high pressure via feed conduit
401 for injection into vertical reactor vessel 41~ above a feed
distributor grid 412, which provides for distribution of a lift gas
passing via conduit 414 throu~h the small diameter holes in the
grid 412. Fluidization is effected in the bottom portion of the bed
by upwardly flowin~ lift gas introduced via conduit 414. Although
depicted without baffles, the vertical reaction zone can contain
open end tubes above the grid for maintaining hydraulic constraints,
as in US Pat. 4,251,484. Optionally, a variety of horizontal
baffles may be added to limit axial mixing in the reactor.
m e~modynamic conditions in the reaction vessel can be controlled by
adjusting feed tei~perature, catalyst temperature and rate, or by
heat exchange means 16. In the reactor configuration shown the heat
exchanger tubes limit mixing in the reactor.
Provision is made for withdeawing catalyst from above grid
412 by conduit means 417 pfovided with flow control valve means to
control passage via air lift line 418 to the catalyst regeneration
system in vessel 420 where coked catalyst particles are oxidatively
regenerated in contact with aif o~ other regeneration gas at high
tempeeature. In order to add sufficient heat to the catalytic
reaction zone 410, eneegy may be added by combustion of flue gas or
othe~ fuel steeam in the regenerator Regenerated catalyst is
retufned to the reactor fluid bed 410 through conduit means 422
peovided with flow contfol valve means. The hot regenerated
catalyst is charged to the catalyst bed sufficiently below the upper
nterface to achieve good mixing in the fluid bed. The rate of flow
for eegeneeated catalyst may be adjusted to provide the degree of
theemal input fequied fof effecting endothermic conversion, and the
fate will depend upon the amount and composition of the alkane
components
: ::
: ~ :
,
~,91~9
4178~ -13-
Initial fluidization is achieved by forcing a lift qas
upwardly through the catalyst. A light gas, with or without diluent
or recycle, may be charged at a bottom portion of the reaCtQr
beneath grid 412. Pressurized liquid feedstock is introduced above
reactant distrihutor grid 412, and pumped to one or re spray
nozzle means. The liquid is dispersed into the bed of catalyst
thereabove at a velocity sufficient to form a generally upwardly
flowing suspension of atomized liquid reactant with the catalyst
particles and lift gas.
Cyclone separator means may be positioned in an upper
portion of the reactor vessel. m e product effluent separated from
catalyst particles in the cyclone separatinq system then passes to
effluent separation system 430. The product effluent is cooled and
separated to recover C5+ liquid hydrocarbons, gaseous recycle or
offgas, along with any byproduct water or catalyst fines carried
over. A portion of the light gas effluent fraction may be recycled
by compressing to form a motive gas for the liquid feed or recycle
for use as lift gas. The recovered hydrocarbon product comprising
C5 olefins and/or aromatics, paraffins and naphthenes is
thereafter processed as described hereafter to provide a desired
aromatic product.
In Fig. 5, directed to maximum recovery of BTX, the
feedstock stream 501 is injected into reactor vessel 510 containing
the fluidized bed of catalyst, along with fluidizing gas stream 514
and recycle streams 516, 518. Reactor effluent is cooled in heat
exchanger 520 and partially separated in a series of phase
separation drums, HTS 522 (high temperatuee separator) and LTS 524
(low temperature separator). A light gas stream may be recovered
from LTS 524, pressurized in compressor 528, and recycled via
conduit 5I4 to comprise at least a portion of the lift gas.
Condensed liquid from the separators 522, 5~4 is fed to a
debutanizer tower 530, along with a portion of the I.TS overhead
vapor. Debutanizer overhead vapor is further fractionated by
deethanizer tower 532 from which offgas stream 534 is recovered.
. . ~ ~, .,, i
~.917fi9
417~ 14-
m is liqht hydrocarbon stream may be employed as fuel gas in the
regenerator vessel. Deethanizer liquid bottoms, rich in C3-C4 LPG
alkanes, may be recovered via line 536 as product or recycled for
further conversion via conduit 518 to reactor 510. m e C5+ liquid
from the debutanizer 530 is passed to a liquid-liquid extraction
unit 540 for recovery of the aromatics components with a selective
solvent, such as sulfolane, etc. Following fractionation of the
solvent phase, an aromatics product stream 542 is recovered,
containing at least 50~ BTX. The C5~ aliphatics components may be
recovered as a product gasoline stream; however, it is advantageous
to recycle this stream for further conversion to increase the net
aromatic product.
Under optimized process conditions the turbulent bed has a
superficial vapor velocity of about 0.3 to 2 meters per second
(m/sec). At hi~her velocities entrainment of fine particles may
become excessive and beyond lO 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. Fven 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, preferrably 300 to 500, 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. m is density is generally between the catalyst
concentration employed in dense beds and the dispersed transport
systems. 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 a clean
apparent density of 1.06 gm/cc and packed density of 0.85, an
average fluidized bed density of about 300 to 500 kg/m3 is
satisfactory.
~?~91~
4178+ -15-
By virtue of the turbulence experienced in the turbulent
regime, gas-solid contact in the catalytic reactor is improvedr
providing substantially complete conversion, enhanced selectivity
and temperature uniformity. One main advantage of this technique is
the inherent control of bubble size and characteristic bubble
lifetime. Bubbles of the gaseous reaction mixture are small, random
and short-lived, thus resulting in goo~ contact between the gaseous
reactants and the solid catalyst particles.
A significant difference between the process of this
invention and conversion processes of the prior art is that
operation in the turbulent fluidization regime is optimized to
produce high octane C5 liquid in ~ood yield. The weight
hourly space velocity and uniform contact provides a clo~se control
of contact time between vapor and solid phases, typically about 3 to
25 seconds. Another advantage of operating in such a mode is the
control of bubble size and life span, thus avoiding large scale gas
by-passing in the reactor. The process of the present invention
does not rely on internal baffles in the reactor for the purpose of
bubble size control such as the baffles which are employed in the
prior art dense bed processes discussed above.
As the superficial gas velocity is increased in the dense
bed, eventually slugging conditions occur and with a further
increase in the superficial gas velocity the slug flow breaks down
into a turbulent fegime. The transition velocity at which this
turbulent regime occurs appears to decrease with particle size. The
turbulent regime extends from the transition velocity to the
so-called transport velocity, as described by Avidan et al in U.S.
Patent 4,547,616 and by Tabak et al. in U.S. Patent 4,579,999. As
the transport velocity is approached, there is a sharp increase in
the rate of particle carryover, and in the absence of solid recycle,
the bed could empty quickly.
Several useful parameters contribute to fluidization in the
turbulent regime in accordance with the process of the present
invention. When ~mploying a ZSM-5 type zeolite catalyst in fine
X
417S~ ~,9~9
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
apparent "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. The velocity specified here is for an
operation at a total reactor pressure of about 100 to 300 kPa.
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.
The reactor can assume any technically feasible
configuration, but several important criteria should be considered.
The bed of catalyst in the reactor can be at least about 5-20 meters
in height, preferably about 7 meters. Fine particles may be
included in the bed, especially due to attrition, and the fines may
be entrained in the product gas stream. A typical turbulent bed may
have a catalyst carryover rate up to about 1.5 times the reaction
zone inventory per hour. If the fraction of fines becomes larqe, a
poetion of the carryover can be eemoved from the system and replaced
by larger particl~es. It is feasible to have a fine particle
separator, such as a cyclone disposed within the reactor shell to
recover catalyst carryover and return this fraction continuously to
the bottom of the reaction zone for recirculation at a rate of about
one catalyst inventory per hour. Optionally, fine Particles carried
from the reactor vessel entrained with effluent gas can be recovered
by a high operating temperature sintered metal filter.
This process can be used with any process stream which
contains sufficient liquid olefins and dienes. Preferably the feed
is substantially free of deleterious oxygenates and sulfur
X
- 16-
~?.9~7~.~
4178+ -17-
compounds. Because the catalyst can be readily removed and
regellerated, or replaced, the process can tolerate a lot of
impurities in the feed. Experimental runs are performed using a
ZSM-5 catalyst to demonstrate the inventive process. The fluidized
bed unit can be operated over a wide range of process variables and
catalyst activity.
Reactoe Operation
A typical single pass reactor unit employs a
temperature-controlled catalyst zone with indirect heat exchange
and/or adjustable gas quench, whereby the reaction exotherm can be
carefully controlled to prevent excessive temperature above the
usual operating range of about 315C to 650C. Preferably average
reactor temperature is 340C to 430C to maximize production of
gasoline boiling range hydrocarbons. Temperatures of 425-580C
maximize production of aromatics. Heat exchanging hot reactor
effluent with feedstock and/or recycle streams will save energy.
Gptional heat exchangers may recover heat from the effluent stream
prior to fractionation. It is preferred to operate the olefin
conversion reactors at moderate pressure of about 100 to 3000 kPa
(atmospheric to about 400 psig) to maximize liquid yields. Higher
pressures, to 6000 kPa, favoe aromatics yields.
The weight hourly space velocity (WH5V, based on total
olefins in the fresh feedstock is about 0.l-5 WHSV. Typical product
fractionation systems are desceibed in U.S. Patents 4,456,779 and
4,504,693 (Cwen, et al.~.
To prevent premature non-catalytic reaction of the dienes,
it is desirable to maintain reactant li~uid feedstream temperatu~e
below about 180C (350F) until injection into the fluidized bed.
Appropriate thermal insulation or quenching of the feedstream to the
injection point can laegely prevent gum and coke formation in the
liquid phase prior to catalysis.
,
.. .. .
~?.917fi~
4178+ -18-
~ tomization of the pressurized liquid reactant feedstream
can be achieved by known techniques, such as liquid speay nozzles,
motive gas, ultra sonics, etc. A suitable nozzle is shown in Fig.
2, ~herein a concentric feed liquid projection device lO0 is
depicted in vertical cross section view. Pressurized liquid 1OWS
through a supply conduit 123. The nozzle is mounted onto the vessel
internal steucture by screw cap means 130 or similar attachment
means. A motive Eluid supplied under pressure through conduit 12
drives the pressurized liquid flowing from the nozzle orifice 140
for injection into the reaction vessel at sufficient velocity to
induce a fine vertically directed spray of atomized liquid having an
average particle size up to about 300, preferably about 50 microns.
The number and arrangement of nozzles will depend upon the cross
sectional area of the fluidized bed and fluidization characteristics
of the gas-solid-liquid mixture. The atomized stream from a
pressurized nozzle can be made to effect penetration into the bed at
a depth and/or lateral radius of a meter or more. The mixture fluid
may be an inert material, nitrogen, lower aliphatic gas, stream, etc.
Thermal insulation of the liquid diene-containing
feedstream from the hot reaction medium in the reaction vessel can
be achieved by applying to the liquid feed conduit a layer of
thermal insulation, such as a ceramic shield or the like. Jacketed
conduits with heat adsorbing fluid may also be suitable.
EXAMPLE 1
In the present example a C4+ liquid stream is
converted to aromatics-rich gasoline in the fluidized bed reactor
employing acid ZS~1-5 powder catalyst having a fresh alpha value of
about 80 at an average conversion temperature of about 425C (800F)
and total pressure of about 275 kPa (25 psig).
The liquid pyrolysis gasoline feedstock contains about 22
wt. % C4 mono-alkenes, 27% C4+ dienes (mainly 1,3-
butadiene), 49% C4 paraffins, 2% aromatics and naphthenes, and
less than 1% C3 aliphatics.
X
~X917fi~
4178+ -19-
Typical olefinic pyrolysis byproduct streams are shown in
Table 1.
Following initial heating and fluidization of the powdered
catalyst with a heated lift gas (e.g. C2 hydrocarbon), the
feedstream is preheated and maintained below 180C prior to
injection into the bed. After achieving steady state operation at a
reaction severity index tR.I.) of about 1, the effluent conversion
product (less any lift gas components) comprises 82 wt. % C5~
liquid gasoline havin~ a research octane rating of 94 (RON). The
total aromatics content is 18 wt. %, including 1% benzene (B), 5%
toluene (T), 6% xylenes (X) and ethyl benzene, 4~ Cg aromatics
isomers and 10% C10 isomers, mainly durene. The predominant
nonaromatic fraction (65~) contains mainly mono-olefins, paraffins
and naphthenes, and the light gas C4 fraction is 17% of the
conversion product.
~?.917fi9
4178+ -20-
TALLE 1
Example of Diene-Rich Feedstock (ethane cracker byproduct)
Component Vol.
C3 1.0
i-butene 0.08
1,3-butadiene 0.51
t.2,butene 0.1
c.2,butene 0.15
1,2 butadiene 0.14
3m 1 butene 0.45
isopentane 5-44
1,4 pentadiene 0.6
l-pentene 0.63
n-pentane 1.92
isoprene 2.3
c,2,pentene 0.35
2m2butene 0 45
t,l,3, pentadiene 1.5
c,l,3,pentadiene 1.0
20 cyclopentadiene 13.7
cyclopentene 1.7
2,3 d.m. butane 1.7
3moentene 0.85
hexane 0.95
unknown C6 1.04
cyclohexane 3.06
benzene 34 4
unknown C8 3-47
Toluene 10.1
vinyleydohexene 0.19
ethylbenzene 1.29
xylene 1.01
styrene 0 3
unknown Cg+ ~ 6.9
~L?291~fi~
4178+ -21-
The above diene-rich stream example contains C6
aromatic hydrocarbons which can be separated before feeding to the
reactor. Typical ranges of diene-rich pyrolysis gasoline streans
comprised of mainly C4-C6 hydrocarbons are:
Vol. %
Cienes 5-60
Mono-alkenes 5-30
Aromatics 1-5*
Alkanes 20-60
Naphthenes 1-5
*can be as high as 60% if C6~ fraction is not separated.
.
1?~917fi~
417~+ -2~-
EXAMPLES 2-4
A series of continuous olefin conversion runs are conducted
using H-2SM-5 (65%) catalyst having an alpha value of about 175 at
the beginning of the aging runs made under oligomerization
conditions without regeneration to upgrade mixtures of ethene,
propene and butadiene and to determine the effects of diene
concentration on catalyst aging. The control feedstock (Example 2)
is compared with diene-containing feeds in Table 1.
TABLE 2
Example 2 Example 3 Example 4
Ethene 0 0.7 1.8
Propene 26.8 28.1 22.9
Butenes 35.7 31.9 31.7
1,3 Butadiene O (control) û.8 5.1
Alkanes (C4 ) 37.5 38.5 38.5
Recycle (mol/mol olefin) 2.5:1 2.5:1 2.5:1
~X91~6~
4178+ -23-
The conversion unit is a single bed isothermal reactor
using HZSM-S having a crystal size less than 0.5 microns, together
with 35% alumina binder and having a fresh alpha value of 175. The
continuous runs are conducted at about 6600 kPa and weight hourly
space velocity (wHSV) of about 0.8 parts olefin feed per part by
weight of catalyst per hour. The conversion runs are started at
205C (400F) and the temperature is increased to compensate for
coke deposition, while maintaining total olefin conversion of at
least 80~, preferably over 90X. Results of the aging studies are
plotted in Fig. 3, with all conversion rates being normalized to 80%
to 166C+ (330F+) product for comparison purposes. Selectivity of
the conversion product to heavier hydrocarbons is shown in Table 3.
TABLE 3
Example 2 Example 3 Example 4
Total Liquid Product,
50~ pt, C (F) 261/(501) 259/(498) 244/(472)
~istillate Species (As Cut)
5 wt. ~, C (F) 232/(434) 250/(483) 297(477)
95 wt. % C (F) 369/(697) 383/(722) 379(715)
Gravity, API 44.3 41.2 38.9
Aniline Point 177 184 172
.~ .
. . .
~9i~
4178-~ -24-
While the aromatics product content of the control runs
averaged about 2-5~, the 5.1% butadiene feed (Example 4) is upgraded
to an aromatics content of 15.5 wt. %, more than 3 times the diene
input. The average paraffin content is less than 14% and the liquid
dominant product is 70% + olefins and naphthenes.
m ese results indicate butadiene, at levels of 1 wt. %
or less, do not cause siqnificantly increased catalyst aginq or
lower peoduct selectivity. It was surprising that adding diene
would increase the a~omatics yield, and surprisinq that such larqe
aromatics yields could be achieved at reactor temperatures just
above 200C. Typical FCC C3/C4 olefins from a depropanizer feed
stream contain n.3-0.6 wt. % butadiene which is less than the 0.8
wt. % butadiene concentration that was used in this study. Even at
the 5.1 wt. % butadiene level, though catalyst aging was increased,
product selectivity to heavier hydrocarbons r~mained relatively high.
EXAMPLE 5
In the present high severity example a C4 liquid
stream is converted to aromatics-rich gasoline in the fluidized bed
reactor employing acid ZSM-5 powder catalyst having a fresh alpha
value of about 175 at an average conversion temperature of about
480C (900F) and total pressure of about 275 kPa (25 psiq).
The liquid feed was the same as used in example 1.
Following initial heating and fluidization of the powde~ed
catalyst with a heated lift gas (e.g. C2 hydrocarbon), the
feedstream is preheated and maintained below 180C prior to
injection into the bed. After achieving steady state operation at a
reaction severity index (R.I.) of about 2, the effluent conversion
product (less any lift gas components) has an aromatics content of
34.4 wt. %, including 4.9% benzene (B), 11.8% toluene (T), 14~
xylenes (X) and 0.9% ethyl benzene, 2.3% Cg aromatics isomers and
0.5% C10 isomers. The nonaromatic fraction contains mainly
t ?.91~6~?
4178+ -25-
mono-olefins, paraffins and naphthenes, and the light gas C4
fraction is 13.5% of the conversion product.
Comparative effluent streams for high severity and low
severity conversion runs under steady state reactor conditions are
shown in Table 4. High severity operation means a reactor
temperature of 600C. Low severity operation was run at 425C
reactor temperature, at a RI of 2. ~oth were run at 0.87 WHSV, 1
bar (100 kPa) over HZSM-5.
~?.,9176~t
4178+ -25-
TABLE 4
Products EXAMPLE 5A EXAMPLE 5B
Yi High Severity Low Severitv
H2, wt. ~ 3.0 1.3
Cl 13.5 2.6
C2 10.3 6.1
C3 3.8 3.4
c4 2.0 1.4
C5 Non-Aromatic - 50.3
Benzene 22.1 4.9
Toluene 21.9 11.8
Ethyl Benzene 0.9 0.9
Xylene 15.6 14.0
C9 Aromatics 2.7 2.3
C10 3.2 0.5
Coke 1.0 0.5
100. 0 100. 0
~.917~i~
4178~ -27-
The flexibility of the fluid bed operating parameters for
controlling the reactor temperature under exothermic reaction
conditions allows an easy adjustment for achieving the optimal yield
structure.
Fither C5~ liquid yield, or BTX yield may be optimized
when temperatures of 315-650C are used in the reactor. The net
yield o~ aromatics can comprise over 3C~o of the olefins in the
feed. The propane:propene ratio will usually be 0.7:1 to 5:1 to
maximize aromatics.
To fluidize the catalyst at the bottom of the reactor prior
to injection of the liquid feed stream, a lift gas may be employed.
This can be an inert diluent or recyled light gas, such as methane,
ethane, ethene, propane, etc. Recycle of C3 light hydrocarbons
may also be desirable under certain circumstances, for instance with
unreacted aliphatics which requiure further conversion or for
dilution of highly exothermic feedstocks.
The thermodynamic balance of exothermic olefin
oligomerization and endothermic paraffin reactions can have
significant impact on the reaction severity conditions.
The use of a fluid-bed reactor in this process offers
several advantages over a fixed-bed reactor. Due to continuous
catalyst regeneration, fluid-bed reactor operation need not be
adversely affected by oxygenate, sulfur and/or nitrogen containing
contaminants present in the pyrrolysis byproduct.
