Versatile fluidized bed reactor

A laboratory scale fluid catalytic cracking apparatus and method of use thereof, which provides cracking performance that emulates commercial riser cracking. The apparatus includes a reactor having a removable feed injector to quickly facilitate changing hydrocarbon contact time without varying the feed rate or diluent rates, or catalyst charge, and also without the expense of a new reactor. The feed injector is a tube within a tube design. The feed injector allows hydrocarbon feed as well as fluidization gas to be delivered to a prescribed axial position within a catalyst bed in the reactor to directly affect hydrocarbon contact time. The reactor also includes a conical bottom head having a conical section and a lower fluidization gas nozzle connected at its apex. The total included angle of the conical section may vary between 10.degree. and 1700.degree., however it is preferred to be in the range of The reactor geometry combined with the location of the fluid and gas sources generate the desired torroidal motion of the catalyst bed, which is significantly enhanced over conventional designs which do not use multiple nozzles and the conical bottom design.

TECHNICAL FIELD 
The present invention relates to the design of a fluidized bed reactor one 
use of which is for laboratory evaluation of the fluid catalytic cracking 
process with particular regard to catalysts, feedstocks, and process 
parameters. Fluid catalytic cracking is the dominant catalytic process for 
producing transportation fuels and chemical feedstocks worldwide. 
Consequently, extensive efforts have been made at developing useful 
laboratory tests pertinent to this process for the purposes of developing 
improved catalysts, quantifying and correlating the cracking character of 
various feedstocks based on their respective properties, understanding the 
implications of different process conditions, and improving commercial 
process design. The present invention also relates to the design of 
fluidized bed reactors, which have applications in the hydrocarbon and 
chemical process industries on both commercial and laboratory scales. 
BACKGROUND 
The two broad approaches commonly used in the laboratory for studying the 
fluid catalytic cracking process are continuous processing units and batch 
processing units. The continuous processing units are basically 
scaled-down versions of commercial operating units and are typically very 
complex systems that are expensive to construct, operate, and maintain. In 
addition, they require large samples of catalyst as well as feed compared 
to batch laboratory cracking units. Batch processing units use a single 
charge of catalyst (typically less than 200 gms.) and process a small 
sample mass of feed that is usually injected over the catalyst for a 
period of time on the order of a minute. The ratio of the catalyst mass to 
the feed mass is referred to as the catalyst-to-oil ratio and typically 
ranges from 3 to 10. Batch processes provide considerable cost and speed 
advantages over continuous units for laboratory studies, primarily because 
of their relative simplicity and the smaller scale. 
Two types of batch processes are commonly used: fixed bed and fluidized 
bed. Fixed bed type units are not appropriate for processing high boiling 
range feeds, which are often the feed for commercial fluid catalytic 
cracking units (FCCU). For this and other reasons, many laboratories have 
abandoned the use of fixed bed reactors for evaluating commercial FCCU 
operation. Thus, fluidized bed reactors are preferred over fixed bed 
reactors for studying the fluid catalytic cracking process on a laboratory 
scale. 
One of the most important parameters in fluid catalytic cracking is the 
time that hydrocarbons are in contact with catalyst. Research over the 
past several years uncovered that as much as 90% of the feed conversion 
takes place in the short contact time condition in the feed and catalyst 
mix-zone of the riser reactor in commercial FCCUs. This knowledge led to 
the revamp of older and larger reactors to smaller designs, because 
contact time dramatically affects yields and performance. Normally, 
reducing the contact time requires that the commercial unit operate at a 
higher catalyst-to-oil ratio than when the contact time is longer. It is 
therefore important for laboratory scale fluidized catalytic cracking 
apparatus to provide the flexibility to vary contact time and 
simultaneously operate at high catalyst-to-oil ratios. 
There are several ways to vary contact time in laboratory fluid-bed 
reactors. The widely known techniques of altering the hydrocarbon feed 
rate, the rates of any diluent gases, and/or altering the catalyst charge 
provide results, however, they are not entirely consistent with commercial 
experience or have other deficiencies which limit their applicability. 
Walsh, U.S. Pat. No. 4,419,328, discloses a laboratory apparatus for 
investigating the performance of catalytic cracking catalyst utilizing 
batch techniques and a fluidized bed reactor. Walsh teaches the use of the 
laboratory apparatus and techniques for obtaining cracking data but not a 
reactor apparatus or method that can be utilized to emulate the 
performance characteristics expected in a commercial scale reactor. Walsh 
discloses a fluidized bed reactor but it does not include a movable feed 
injector or disclose in any way the important aspects pertaining to 
injector location and its relationship to controlling hydrocarbon contact 
time. Walsh does not disclose or teach the injection of multiple feeds at 
different locations. In addition, Walsh does not disclose or teach an 
apparatus or method to achieve the catalyst circulation pattern within the 
reactor including its relation to commercial catalytic cracking as well as 
ways to enhance the circulation by proper reactor design and injector 
positioning. 
It is well known that fluidized bed reactors of many different designs 
often exhibit a preferred flow of solids up the center core and return of 
solids down the normally cylindrical containing wall. Perry's Chemical 
Engineers' Handbook (Copyright 1984) mentions this tendency (Section 20 
pg. 66) and the effect is noted in U.S. Pat. No. 5,262,104-Schwartz. In 
these references, however, there is no recognition of the potential 
applications of the circulation pattern nor are any ways to enhance the 
circulation provided. 
SUMMARY 
The present invention provides a method and apparatus for contacting 
hydrocarbon feed and catalyst in a laboratory or commercial scale 
fluidized bed and also controlling the time the hydrocarbons are in 
contact with the catalyst. More specifically, the axial position of the 
feed injector may be accurately positioned to vary the distance traveled 
by reactant and product hydrocarbons before they exit the catalyst bed. 
The closer the feed point is to the top or exit of the fluid bed the 
shorter the contact time between hydrocarbon and catalyst. This technique 
offers several advantages over the other known methods for varying contact 
time: altering the hydrocarbon feed rate, diluent rates, and/or the 
catalyst mass. 
If the feed rate is increased to reduce the hydrocarbon contact time then 
the catalyst circulation pattern changes in an uncontrollable way as does 
the contacting of the feed and catalyst. In addition, at constant 
catalyst-to-oil ratio, varying the feed injection rate requires changing 
the injection time, which adds complication to the kinetic analysis and 
comparison of the results with commercial scale FCCUs. The present 
invention shows that moving the feed injector does not significantly alter 
the catalyst fluidization pattern, as compared with altering the feed 
rate, and does not require any changes in the injection time to attain a 
prescribed catalyst-to-oil ratio. 
Varying the diluent rates has the same effect on catalyst patterns as 
changing the feed rate. In addition, varying diluent rates changes the 
hydrocarbon partial pressure which affects yields and performance. Simply 
moving the feed injector does not affect hydrocarbon partial pressure. 
Decreasing the catalyst load to decrease contact time reduces bed height 
and can result in a short bed within which it is more difficult to 
distribute the hydrocarbon feed and produces more erratic yields. In 
addition, as catalyst load is reduced it becomes difficult to measure 
yields at high catalyst-to-oil ratio which is inherently important to 
reduced contact time operation. This difficulty is because the feed mass 
is very low and difficult to account for in a mass balance (at high 
catalyst-to-oil ratio with reduced catalyst mass). In addition, at 
constant catalyst-to-oil ratio, varying the catalyst load to affect 
contact time also requires changing the feed injection time, which adds 
complication to the kinetic analysis and comparison of results with 
commercial scale FCCUs. 
Another feature of the invention couples the reactor geometry with the gas 
and feed supply rates and injector locations to provide enhanced catalyst 
circulation up the center core of the reactor and down the walls. This 
pattern is natural to fluidized beds, but it is enhanced by the present 
invention and such enhancement improves the performance so as to closely 
replicate that of a commercial scale FCCU. The conical bottom eliminates 
catalyst dead zones or regions of poor circulation and the feed injector 
supplies feed directly onto the catalyst. The catalyst circulation 
provides a high local ratio of catalyst-to-feed at the tip of the feed 
injectors at any point in time. This results in a more complete 
utilization of the catalyst. The enhanced catalyst circulation also 
provides for improved heat transfer to and from the catalyst and reactor 
contents at the wall surface of the reactor. Many applications may benefit 
from improved heat transfer. 
One objective of the present invention is to provide a laboratory reactor 
design for evaluating the fluid catalytic cracking process that yields 
results that are comparable to commercial operations including a method 
for studying the effects of varying the hydrocarbon contact time. 
Another objective of the present invention is to provide a reactor design 
geometry that yields a highly coherent catalyst circulation pattern, which 
is well suited for providing a performance closer to that of an ideal plug 
flow reactor rather than a continuous stirred tank reactor with respect to 
both the contacting of hydrocarbon and catalyst, and the conversion of the 
feed materials to products. This aspect of the present invention enables 
it to be utilized in many applications in the hydrocarbon processing and 
chemical processing industries. The fluidized bed resulting from the 
present invention has considerably more stability, more repeatable 
hydrodynamics, and more efficient catalyst utilization. These enhancements 
provide more precise and relevant performance over a wider range of 
velocity and may allow for higher feed throughput without degradation in 
performance. 
Another objective of the present invention is to provide a versatile 
fluidized bed reactor that emulates the yield performance of commercial 
fluid catalytic cracking units, which inject feed to be cracked at one or 
more feed locations in a nearly ideal plug flow riser reactor. Heretofore, 
either different equipment or a totally new reactor was necessary to 
accomplish the desired testing in the laboratory. Multiple feed locations 
are necessary since different feeds require different contact times to 
obtain optimal yields. The present invention accomplishes this in that 
contact time may be varied by the methods described herein utilizing the 
reactor of the present invention without modification to accommodate 
multiple feeds while producing commercially relevant results. 
Another objective of the present invention is to provide a reactor design, 
which may be used in processes other than fluid catalytic cracking. For 
example, the methods and apparatus described herein may be used in 
catalyst deactivation of fluid catalytic cracking catalyst as well as for 
chemical reactions including but not limited to partial oxidation 
reactions, an example of which is oxidation of ethylene to produce 
ethylene oxide.

DETAILED DESCRIPTION OF THE INVENTION 
FIG. 1 is a schematic representation of a fluidized bed reactor of the 
present invention. The fluidized bed reactor 20 shown in FIG. 1 comprises 
a reactor shell 21 having a top head 22 and a bottom head 23 connected at 
either end of the shell forming a pressure vessel. 
A feed injector 24 is shown extending from above the reactor top head 22 
axially through the top head 22 and into the interior of the reactor 20, 
and a catalyst bed 25. The preferred embodiment of the feed injector 24 is 
a tube within a tube configuration as illustrated in FIG. 1. The inner 
tube of the feed injector 24 includes a hydrocarbon and carrier gas feed 
line 1, which extends axially through an upper fluidization gas line 2 
comprising the outer tube of the feed injector 24. While the feed injector 
24 is shown extending through the reactor top head 22, it may be 
positioned at any location so as to allow the feed injection point to be 
located at any point within the catalyst bed 25. The feed injector 24 is 
removably connected to the reactor top head 22 via injector coupling 3. 
The injector coupling 3 (as well as the feed line coupling 6 and bottom 
head coupling 11 discussed below) comprises a SWAGELOK.RTM. or similar 
type fitting means. A compression fitting 4 or similar means is disposed 
about the exterior of the fluidization gas line 2 to seal against a seat 
within the reactor top head 22 such that when the injector coupling 3 is 
connected to the reactor top head 22 a pressure tight seal is formed. One 
who is skilled in the art will understand the injection point, designated 
by upper arrow 26, of the fluidization gas as well as the hydrocarbon feed 
and carrier gas within the catalyst bed 25 is determined by the length of 
tubing comprising the injector 24, which extends below the compression 
fitting 4 seated within the reactor top head 22. Therefore, to change the 
injection point within the catalyst bed 25, the injector coupling 3 is 
simply removed from the reactor top head 22, the respective feed injector 
is removed and a new injector of a prescribed length is installed in its 
place. The feed line 1 similarly has a compression fitting 5 disposed 
about the exterior of the feed line 1 and seals against a seat within the 
upper end of the fluidization gas line 2. The feed line 1 is removably 
retained by feed line coupling 6 together with the compression fitting 5 
forming a pressure tight seal. While the outlet or injection point of both 
the feed line 1 and the fluidization gas line 2 as shown in FIG. 1 
terminate at the same point within a catalyst bed 25, it is within the 
scope of the present invention that each may be terminated at different 
points within the catalyst bed 25. The feed injector 24 and fluidization 
gas line 2 are used to supply inert gas, such as nitrogen or steam, to the 
feed injection point, upper arrow 26, to both prevent catalyst particles 
from entering and plugging the feed line 1 and to cool the injector. The 
feed injector 24 of FIG. 1 is shown as a tube within a tube directed 
downward into the interior catalyst bed 25 of the reactor 20. However, the 
feed injector 24 may as noted above, be connected through the bottom 
aiming vertically upward. An injector centering means 27 may be connected 
directly to the bottom surface of the reactor top head 22 or to the 
reactor shell 21. The centering means 27 provides a guide path for the 
feed injector 24 so that the injector 24 remains aligned with the axis of 
reactor 20. 
The bottom head 23 includes a conical section 7 of the reactor 20. The 
bottom head 23 also includes a fluidization nozzle 8, connected at the 
bottom or apex of the conical section 7, through which a lower 
fluidization gas line 9 extends upward to provide fluidization gas, such 
as nitrogen or other diluent, to the catalyst bed 25, and if desired 
serves as an alternate source for feed injection. The total included angle 
of the conical section 7 designated as .theta. may vary from 10.degree. to 
170.degree., but in the preferred embodiment the total included angle of 
the conical section is in the range of 20.degree. to 60.degree.. The 
fluidization gas line 9 extends through the interior of the fluidization 
nozzle 8 and terminates approximately at the apex of conical section 7. As 
similarly described above, a compression fitting 10 is disposed about the 
exterior of the fluidization line 9 to sealingly engage with a seat within 
the fluidization nozzle 8. A fluidization coupling or bottom head coupling 
11 is removably connected to the fluidization nozzle 8 to provide a 
pressure tight seal between the fluidization gas line 9 and the interior 
of the reactor 20. As discussed above, a second feed injector 24 as well 
as a second fluidization gas line may be connected to the fluidization 
nozzle 8 in place of fluidization line 9 and each may extend to any 
desired level within the catalyst bed 25. 
A catalyst nozzle 12 is shown at an angle connected at one end to the wall 
of the reactor shell 21 in the region above the catalyst bed 25, and 
isolated with nozzle cap 13. The catalyst nozzle 12 serves as the conduit 
through which catalyst may be loaded into or catalyst retrieved from the 
reactor 20 by conventional means well known in the art. The catalyst 
nozzle 12 may be connected using a compression fitting and line as 
detailed above to provide flow to both catalyst addition and retrieval 
hardware (not shown). 
The inert gases as well as the product gas are removed from the reactor 20 
through effluent product nozzle 14. Effluent product line 15 extends 
axially through the effluent product nozzle 14 providing a conduit for the 
flow of gases to other laboratory equipment (not shown). A compression 
fitting 16 is disposed about the exterior surface of the effluent product 
line 15 and seals against a seat within the outlet of the effluent product 
nozzle 14. A pressure tight seal is created by connecting effluent 
coupling 17 to the outlet of the effluent product nozzle. A product filter 
18 is connected at one end of the effluent product line 15 to prevent 
carry over of catalyst with the gases. 
To prepare the reactor 20 for commercial simulation testing of a FCCU 
process, the appropriate feed injector 24 is first inserted through the 
top head 22 and the injector coupling 3 is connected to the top head 22 to 
establish a fluid tight seal. The feed injection point (designated by 
arrow 19) is now established with respect to the bottom of the catalyst 
bed, designated by lower arrow 26, and thus the catalyst-to-feed contact 
time is established for the prescribed catalyst charge and injection 
rates. The reactor 20 is then charged with the desired mass of catalyst 
through the catalyst nozzle 12 while fluidization gas is allowed to flow 
through the feed line 1, and the fluidization gas lines 2 and 9 into the 
catalyst bed 25. This results in the desired torrodial motion of the 
catalyst bed 25. The reactor 20, catalyst bed 25, and fluidization gases 
are operated at a desired temperature via temperature control means (not 
shown). The temperature control means may be an external jacket type 
heating element or other means well known in the art. The feed may then be 
injected through the feed line 1 and into the interior of the reactor 20 
so as to come in direct contact with the catalyst bed 25. 
The center or core of the catalyst bed 25 is of lower density since it 
includes gas flows and vapor phase cracked products in addition to 
catalyst. These materials thus flow upward and almost all of the 
hydrocarbon vapor products escape out of the top of the catalyst bed 
exiting through the effluent product line 15, while the catalyst 
circulates primarily back down along the interior wall of the reactor. The 
exiting vapor phase cracked products may be further processed and analyzed 
via conventional laboratory equipment. To begin a series of runs with a 
different hydrocarbon contact time, the catalyst bed 25 is pneumatically 
removed from the reactor 20 via catalyst nozzle 12 to catalyst removal 
means (not shown) by flowing inert gas through the lower fluidization gas 
line 9. The feed injector 24 is then simply removed and a different 
injector is installed. The process described above may now be repeated. 
One skilled in the art will understand that the time the hydrocarbons 
contact the catalyst is directly related to the distance from the top of 
the catalyst bed to the feed injection point. Therefore, by solely varying 
the location of the feed injector 24 relative to the top of the catalyst 
bed 25 by gauging the distance of the injector from the apex of the 
conical section 7 illustrated by lower arrow 26 to a new location, it is 
possible to systematically alter the reaction contact time. 
Techniques and advantages over prior art methods and apparatus are 
illustrated in the following non-limiting examples: 
EXAMPLE 1 
To illustrate the operation, three feed injector locations were used to 
crack the feedstock described in Table 1 over the catalyst described in 
Table 2. The catalyst charge to the reactor is 9.0 gms and the cracking 
temperature is 990.degree. F. (532.degree. C.). The results are compared 
in Table 3 at a catalyst-to-oil ratio of 5.0 and some of the salient 
results are illustrated in FIGS. 2 through 6. 
TABLE 1 
______________________________________ 
FEED PROPERTIES 
Feedstock Feed A 
______________________________________ 
API Gravity 21.2 
Specific Gravity, 60/60.degree. F. 0.927 
Sulfur, wt % 0.87 
Conradson Carbon Residue, wt % 0.8 
Distillation (D 2887) wt % .degree. F./.degree. C. 
10 659/348 
50 833/445 
90 983/528 
______________________________________ 
TABLE 2 
______________________________________ 
CATALYST PROPERTIES 
Catalyst Equilibrium Cat. 
______________________________________ 
Total SA, m.sup.2 /gm 
213 
Zeolitic SA, m.sup.2 /gm 142 
Matrix SA, m.sup.2 /gm 71 
Z/M 2.0 
RE.sub.2 O.sub.3, wt % 2.4 
UCS 24.34 
Nickel, ppmw 1400 
Vanadium, ppmw 2500 
______________________________________ 
TABLE 3 
______________________________________ 
YIELD AT CONSTANT CAT-TO-OIL 
1.125" 2.125" 2.625" 
FROM FROM FROM 
INJECTOR LOCATION BOTTOM BOTTOM BOTTOM 
______________________________________ 
Temperature, .degree. F. 
990 990 990 
Temperature, .degree. C. 532 532 532 
430.degree. F.+ Conversion, wt % 73.06 71.04 65.38 
Catalyst-to-Oil, wt/wt 
#STR1## 
#STR2## 
#STR3## 
Delta Coke, wt % 
1.10 0.98 0.93 
YIELDS, WT %: 
Coke 5.48 4.92 4.67 
Dry Gas 3.11 2.73 2.43 
Propane 1.53 1.26 1.06 
Propylene 4.44 4.16 3.65 
n-Butane 1.24 1.06 0.88 
Isobutane 4.50 3.95 3.40 
C4 Olefins 4.81 4.72 4.41 
Gasoline (C5-430.degree. F.) 47.94 48.25 44.89 
LCO (430-650.degree. F.) 17.39 17.84 18.43 
650.degree. F.+ 9.56 11.11 16.19 
TOTAL 100.00 100.00 100.00 
______________________________________ 
The contact time is varied by moving the feed injector axially. The 
injector positions noted in the Figures are measured from the reactor 
bottom in accordance with FIG. 1. The relative contact time for the 
positions are: 1.125" injector, 1.00 relative contact time; 2.125" 
injector, 0.60 relative contact time; and 2.625" injector, 0.36 relative 
contact time. A relative contact time of 1.00 normally compares to the 
performance of a commercial unit with a riser contact time of 3 to 4 
seconds. For these data sets the catalyst charge, feed injection rate, and 
diluent rates are held constant. 
Decreasing contact time impacts conversion, coke, and delta coke as shown 
in FIGS. 2 through 4. At constant coke, as contact time decreases (going 
from the 1.125" injector to the 2.625" injector), conversion decreases, 
delta coke decreases, and required catalyst-to-oil increases. Hydrocarbon 
contact time influences dry gas and bottoms conversion as shown in FIGS. 5 
and 6. 
Table 3 shows all of the yield shifts characteristic of decreasing contact 
time by comparing the data at constant catalyst-to-oil. Reducing contact 
time of commercial operations usually increases gasoline yield. As very 
short contact times are employed, however, the gasoline yield begins to 
decrease because of conversion loss. This is effectively shown in at 
constant catalyst-to-oil ratio in Table 3. 
EXAMPLE 2 
To illustrate the relevance of the particular laboratory apparatus, data 
are provided in Table 4 comparing the laboratory unit to a particular 
commercial, full-scale fluid catalytic cracking operation with both the 
laboratory and commercial unit operating on the same feed and catalyst. In 
the table the ratio of the laboratory yields to commercial yields are 
indicated and the closer the values are to unity the closer is the 
correspondence of the performance of the lab and commercial unit. 
The commercial unit is a fully modernized, short contact time unit which is 
operating well (with radial feed nozzles, good riser termination, and good 
stripper design). The operation of the Invention involved tuning the 
performance by moving the feed injector until the laboratory 
catalyst-to-oil ratio was within about 10% of the commercial 
catalyst-to-oil ratio (at the commercial level of conversion). The data is 
for a cracking temperature (initial fluid-bed) set at the commercial riser 
outlet temperature. 
The data shows that the current invention performs within 10% of the 
reported commercial yields except C4 saturates. These offsets are simple 
to tune or further calibrate for by adjusting cracking temperature closer 
to the average temperature of the commercial riser. Overall, the 
comparison shows the invention provides data very close to commercial 
operation for modern fluid catalytic cracking units. 
TABLE 4 
______________________________________ 
COMISON OF LABORATORY 
TO COMMERCIAL DATA 
AT CONSTANT CONVERSION 
Laboratory/Commercial 
Parameter Ratio 
______________________________________ 
Cat/Oil 0.88 
Coke 1.08 
Dry Gas 1.05 
LPG 1.07 
Propane 1.04 
Propylene 1.08 
C4 Saturates 1.25 
C4 Olefins 0.92 
Gasoline 0.96 
LCO 0.96 
Bottoms (650.degree. F.+) 1.04 
C3 Saturation 0.97 
C4 Saturation 1.17 
______________________________________ 
EXAMPLE 3 
To illustrate the precision of the laboratory apparatus, which is a direct 
measure of the stability and repeatability of the fluidized bed, yield 
data for repeated runs are provided in Table 5. These data are with the 
same feed and catalyst described in Tables 1 and 2 and a temperature of 
990.degree. F. (532.degree. C.). 
The relative error values of Table 5 indicate the invention and associated 
analytical equipment are very precise. The relative error is nominally 2% 
or less for each yield. Since the analytical equipment used for these 
measurements are no more precise that the data itself, it is clear that 
the invention provides extremely reproducible performance. This is an 
indication of the stability of both the catalyst and feed contacting and 
the catalyst circulation--issues discussed further in Example 4. 
TABLE 5 
__________________________________________________________________________ 
RESULTS FROM REPEAT RUNS AT 5 C/O 
STATISTICAL 
CRACKING DATA SUMMARY SUMMARY 
Run No. 42 46 65 69 MEAN 
STD. 
REL ERR. 
Date 11/8/96 11/8/96 11/26/96 11/26/96 VALUE DEV. (%) 
__________________________________________________________________________ 
Recovery, wt % 
99.2 
99.5 
98.4 99.3 99.1 
0.5 
0.5 
Cet-to-Oil, wt/wt 5.00 5.00 5.00 5.00 5.00 -- -- 
430.degree. F.+ Conv., wt % 72.96 72.93 73.01 72.67 72.89 0.15 0.2 
YIELDS, wt % 
Coke 5.48 5.45 5.60 5.45 5.49 0.08 1.4 
Dry Gas 3.11 3.13 3.08 3.07 3.10 0.03 0.9 
Hydrogen 0.10 0.11 0.09 0.10 0.10 0.01 5.5 
H.sub.2 S 0.46 0.46 0.46 0.46 0.46 0.00 0.00 
Methane 1.11 1.11 1.08 1.06 1.09 0.03 2.4 
Ethane 0.69 0.69 0.70 0.70 0.69 0.00 0.4 
Ethylene 0.76 0.76 0.75 0.76 0.75 0.01 0.8 
Propane 1.59 1.50 1.60 1.52 1.55 0.05 1.1 
Propylene 4.51 4.43 4.44 4.38 4.44 0.05 1.1 
n-Butane 1.29 1.24 1.31 1.25 1.27 0.04 2.8 
Isobutane 4.65 4.43 4.71 4.48 4.57 0.13 2.9 
C4 Olefins 4.76 4.83 4.67 4.72 4.75 0.07 1.4 
1-Butene 1.08 1.08 1.07 1.06 1.07 0.01 0.9 
Isobutylene 1.07 1.11 1.05 1.07 1.08 0.02 2.2 
c-2-Butene 1.12 1.13 1.10 1.11 1.12 0.02 1.5 
t-2-Butene 1.44 1.47 1.42 1.43 1.44 0.02 1.5 
Butediene 0.05 0.04 0.04 0.05 0.05 0.004 7.7 
Gasoline 47.57 47.92 47.59 47.81 47.72 0.17 0.4 
LCO 17.43 17.45 17.49 17.51 17.47 0.04 0.2 
650.degree. F.+ 9.61 9.62 9.50 9.81 9.64 0.13 1.3 
TOTAL 100.00 100.00 100.00 100.00 
__________________________________________________________________________ 
EXAMPLE 4 
To develop the basic design of the invention of FIG. 1, a cold flow (room 
temperature), glass model was constructed and various reactor 
configurations tried. By observing and photographing the catalyst and gas 
interaction, the geometry and gas rates are optimized to provide the 
stable fluidized bed which provides the results of Examples 1 through 3. 
Notes from the comparison of the reactor configurations of FIG. 2 are 
provided in Table 6. 
For the cold flow studies of this example, the superficial velocity is 
varied by increasing the upper fluidization flow rate to both reactors of 
FIG. 2. For Reactor A the lower fluidization flow is set at 180 sccm and 
held constant and the feed injector is 1.125 inches from the reactor 
bottom. The injector diameter is 0.125 inches for Reactor A and 0.250 
inches for Reactor B. 
The comments in Table 6 may be summarized as follows (with due 
qualification for the catalyst material studied): Reactor A (invention) 
maintains a very stable fluidized bed up to 0.50 ft/sec and the stability 
begins to deteriorate at velocities where solids entrainment becomes a 
factor. Reactor B (and many other designs) begins to slug (as large 
bubbles form) at superficial velocities as low as 0.15 to 0.25 ft/sec. 
Reactor A provides stable fluidization over a superficial velocity range 
that is about twice that of Reactor B. The catalyst circulates within 
Reactor A several times per minute as noted in Case 5 of Table 6 (at 0.35 
ft/sec superficial gas velocity). 
TABLE 6 
__________________________________________________________________________ 
VISUAL COMISON OF REACTOR 
CONFIGURATIONS OF FIG. 7 
Bed Height, 
cm 
Superficial Reactor 
Case 
Velocity 
A B Comments 
__________________________________________________________________________ 
1 0.00 cm/sec 
6.0 
6.0 
Catalyst charge to each vessel is 9.0 grams of 
0.00 ft/sec an equilibrium FCC catalyst with 0.82 gms/cc apparent 
bulk density. Average particle size of 79 microns 
2 1.52 cm/sec 8.0 8.0 Catalyst is fluidized. Reactor A is 
circulating 
0.55 ft/sec even at this low velocity. There are catalyst dead 
zones at the bottom of Reactor B. 
3 4.57 cm/sec 8.5 8.5 Some large bubbles are forming in Reactor B. 
0.15 ft/sec Reactor A remains very stable. catalyst 
circulation is 
several times faster in Reactor A than in Reactor B. 
4 7.62 cm/sec 8.7 8.5 Reactor A is stable with apparently no large 
bubbles. 
0.25 ft/sec Slugging is severe in Reactor B. Slugging makes 
measuring bed height difficult (value is an average). 
5 10.7 cm/sec 9.0 9.0 Reactor A is still a dense bed. Catalyst circulate 
s 
0.35 ft/sec several times per minute. Reacotr B experiences 
extreme slugging with ver large bubbles. 
6 18.2 cm/sec 9.0 9.0 Reactor A is still a dense bed along the walls. 
There is 
0.50 ft/sec more bubble aggregation and slight slugging at the 
top. Reactor B is extremeely turbulent/incoherent 
7 21.2 cm/sec -- -- At this velocity catalyst is entrained 
from both reactor 
0.70 ft/sec designs. There is a dense catalyst phase in the lower 
5 cm of Reactor A but catalyst is being entrained. 
__________________________________________________________________________ 
The axial injector in Reactor A was also moved in a series of flow studies. 
As the injector is lowered into the conical bottom the catalyst 
circulation pattern deteriorates and more gas bubbles and slugging occur. 
Apparently, the catalyst circulation pattern is more stable with the 
injector above the conical bottom section. As the injector is elevated 
above the conical section the fluidization pattern remains stable and 
circulates with the toroidal motion shown in FIG. 1 at all higher levels. 
Since the axial position of the injector is directly related to contact 
time, it is clearly possible to supply one feed at one location and a 
second feed at another. In this way, the apparatus could be used to 
perform studies at multiple feed locations pertinent to commercial 
catalytic cracking operations. 
Moreover, while the particular embodiment of the invention has been shown 
specific to fluid catalytic cracking, modifications or application of the 
invention to other catalytic processes by someone skilled in the art of 
reactor design are within the spirit and scope of the invention. Since the 
laboratory apparatus works well at emulating continuous, short contact 
time fluid catalytic cracking operations, it will work well in other 
applications. For example, partial oxidation reactions like ethylene to 
ethylene oxide, and isomerization reactions could be performed in the type 
of fluid bed shown in FIG. 1. In addition, there are many fluidized bed 
applications that may benefit from the invention where the solid is not 
strictly catalyst, the mobile phase is not strictly gaseous, and the 
process may not involve chemical reaction. Use of the invention in these 
applications is also considered within the spirit and scope of the 
invention. 
In addition, also within the spirit and scope of the invention is 
continuous fluidized bed processing. Clearly, it is straightforward to 
apply the batch-wise concepts of this invention to fluid bed reactors 
operating in a continuous mode. 
Near plug flow performance in the central core of the reactor with both 
excellent reproducibility and excellent heat transfer along the reactor 
wall is clearly achieved when the catalyst circulation pattern is enhanced 
by proper combination of geometry, gas flow rates, and injector location. 
While the invention has been described in detail and with reference to 
specific embodiments thereof, it will be apparent to one skilled in the 
art that various changes and modifications can be made therein without 
departing from the spirit and scope thereof. It is to be understood that 
all matter herein set forth or shown in the accompanying tables and 
figures is to be interpreted as illustrative and not in a limiting sense. 
Accordingly, the foregoing description should be regarded as illustrative 
of the invention whose full scope is measured by the following claims.