Fluid Bed Catalytic Cracking Plant R. P. TRAINER, N . W. ALEXANDER,
AND
FREDERICK KL'NRECTHER
Shell Oil Company, Incorporated, Hoaston, 1e.c.
hH
;,barrels
i d bed cat.alytie cracking pilot plant with a capacity per day llaS bee., ,lsed fc,r piloti,,g er,nrnrercinl plant operation and for the study of prooens variables. T h e physical layout and the process design are described with special emphasis on instrumentation, capacity limitations, and nmchanieal featwres peeutinr tu plants using the fluid bed technique. T h e start-up proocdiire i s dcscribed, a6 are sampling techniques, analytical require meiita. and personnel organirntion.
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iTALYT1C eraeking is B pro , developed aboul tell years ago, for convertiuy petrolcum fractions in t.he iuel oil boiling range into gasoliiic arid d i e r lower toiling hydrocarbons. It, givw a htgher yidd of gasoline of hettt'r qualily than can be produccd by thermal cracking and greater yields of propylenc, isobut.itne, and but,ylensfi which w e important, raw mat,erials kj? syrrt,h&c rubber and IWoct,ane aviation ghyolino. This country's entrance into World War I1 made a rapid cxpension of catalytic crsoking facilities imperative. Of several processes available at that time, the fluid bed process, e m p l o y i n g powdered catalyst, eompmed prineipally of silica and alumins, s h o w e d e o TI s i d e r &b I e promise. Much process d e v e l o p m e n t work remeined t o he done bcfore this type or unit could be elosi?ly designed s n d o p e r a t e d for o p t i m u m yields. The Sliell Oil Cornpany, Inc., had installed during the first years of ihc war several fluid bod cat,alybic cracking units rat,ed at B charge capacity of about 16,000 bsrrels per day. Thr operation and, to some extent, the design of t h e w units were guided by thc data obt.ained from several o i the company's fluid bed pilot plants wit,h capacities of 2 to 3 barrels per day. Studies included:
Effect,of regenerntion v a r i a t h on coke burning: rates Evaluation of available feedstocks. such 8 s gas oils and distillates Cat,alytic retrest.ing of sviat,ion gssoline b m stock for improvement Effect of operating eonditions on the cat,alyst activity rate Problems typical of fluid bed processes, such as cetal trition m t t w and entrainment,, catalyst bed densities, el catalyst type and shape on it.s flow characteristirs, and II for meLquring arid controlling catalyst circulation rates Fact,ors affecting spent catslyst stripping (ramoving en and adsorbed hydrocarbons from the catalyst prior t,o rei iion)
The above may be divided into st,udies of nornmlly c tered process variablca and an engineering investigation L lcms tvpiral of the operation of R continuous fluid bed This piqm is concerned wit,h the fluid bed catalytic ci pilot plsnf of t,hr Houst,on Research I,&aratories, Shell 0 pmy, Inc., which hlts a capacity of 3 barrels per day. I primarily with the inoohsnioal and process design featu strumentation. and ing procedure. DESCRIPTION OF I PLANT
T h e building t 81y t i c crackiiij plant is shown in Fi The steel a n d Tr structure is firepro a 30 X 40 ioot floc and is 40 feet high. plan is shown in Fi All instruments and controls are locate^ cont,rsl h e r d close processing equip separate Sp8ces 81 vided for catalyst, 5 nitrogen cylinders, ofke. C h a r g e S t stored in a mpra: iarm and drum area. All utilities I v i d e d b y t h e P, Two 48-inch exhau h a l e d at the top CB
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VoI. 40, No. 1 Figure 3 is a simplified flow sheet of the pilot plant. Oil is vaporized and superheated t o reactor temperature in lead-pot heaters, picks up regenerated catalyst from the regenerator standpipe, and carries it a t a velocity of 15 to 30 feet per second through a 0.75-inch riser to the bottom of the 6-inch reactor, where the velocity drops to 1 foot per second or less. At this lower velocity the catalyst, which is highly dispersed in the transfer line, forms a high density fluidized bed (30 t o 50 pounds per cubic foot) through which the reacting oil vapors rise, keeping the catalyst powder in turbulent motion. Product vapors leaving the catalyst bed pass overhead through stone filters which free them of e n t r a i n e d c a t a l y s t . Liquid hydrocarbons are condensed and withdrawn from a phase separator. Vapors from the phase ¶tor are compressed to 40 pounds per square inch and are partially condensed t o remove heavier hydrocarbons which are recycled to the low pressure phase separator. The uncondensed vapors are metered and sampled continuously. Coke, one of the products of cracking, is deposited on the catalyst in the reactor and thereby deactivates it. I n order to maintain activity t h e catalyst must be regenerated by burning off coke. Regeneration is carried out in the following manner: A stream of spent catalyst is continuously withdrawn from the bottom of t h e reactor t o an external stripping section where entrained and adsorbed hydrocarbons are stripped out with steam, freed of catalyst in a filter section, and partially condensed. Stripped hydrocarbon liquid and vapors are measured and sampled. Stripping spent catalyst not only increases the recovery of hydrocarbon products but also reduces b y 30 t o 40y0the air requirements in the regenerator.
January 1948
INDUSTRIAL AND ENGINEERING CHEMISTRY
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Since the regenerator capacity often limits the capacity of catal y t i c c r a c k i n g units, stripping efficiency is an important process variable. This external stripper has the following advantages over t h e c o m m o n l y used internal stripper: (1) An exact measure of stripping efficiency is obtained, (2) the reactor oil vapors are not diluted with stsipping steam, and (3) the analysis of the isolated stripper products permits the study of stripping mechanisms. The stripper also serves as the spent catalyst standpipe, from which the catalyst is transferred with air through a 1-inch riser a t high velocity t o the 8-inch diameter regenerator. The catalyst temperature is increased from 850" to 950" F. reactor temperature t o 1000" t o llOOo F. as a result of the heat of combustion of coke. In the regenerator bed coke is burned off the catalyst with air and the regenerated catalyst is continuously returned to the reactor. Entrainment of catalyst in the flue gases leaving the regenerator bed is greater than in the reactor because the superficial gas velocity is higher. In order to prevent overloading the regenerated catalyst filters, the g a s e s a r e p a r t i a l l y freed of catalyst in a cyclone separator. Gas from the cyclone passes through the filter section, through a motor valve, and t o the atmosphere. I n contrast t o the reactor, the regenerator filter section is not connected directly t o the regenerator bed. Catalyst has to be withdrawn periodically from the filter housing and returned t o the unit. Continuous product fractionation was not included in the pilot plant, because an unduly large amount of time would be consumed in lining out both cracking and fractionation sections. It was found more satisfactory to take off a total liquid product and subsequently charge it t o laboratory scale batch distillation columns. I n this way, only products from runs of satisfactory control and weight recovery are fractionated. The material balance around the various sections of a fluid bed catalytic cracker is complicated by the fact that the catalyst carries entrained gases from vessel to vessel. This is of considerable consequence in the case of the catalyst flowing from the reactor,
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Vol. 40, No. 1 DRY FLUE GAS,%m
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Figure 4.
CONTAINS
6 % HYOROOEN
Typical Material Balance for Catalytic Cracking Pilot Plant
All numbers in consistent weight units.
Balance excludes instrument inert bleeds and aeration gases
where failure to remove the entrained hydrocarbon vapors results in a n inoreas: in the required regeneration capacity and loss of charge through combustion. I n Figure 4 is shown a material balance taken from a large number of pilot plant runs. The air and steam requirements are based upon 100 weight units of hydrocarbon charge t o the cracking section. The catalyst internal circulation rate is shown t o be constant a t 400 weight units. Of the 100 weight units charged t o the reactor, 94 are recovered overhead from th'e reactor and 6 are carried t o the stripper. Of these 6, 1.5 are stripped off and recovered as useful product and the remaining 4.5 weight units pass t o the regenerator vessel as coke. The catalyst passing from the regenerator always has residual unburned carbon on its surface whichis recirculated into the reactor. The pilot plant is rated at 3 barrels per stream day of West Texas gas oil at about 65% conversion. The upper charge rate is limited by the number of pounds of coke that the regenerator vessel can handle without its temperature exceeding l l O O o F. Above this temperature the catalyst deactivates rapidly. The source of this increased temperature is the 12,000 to 14,000 B.t.u. liberated per pound of carbon burned. It was judged impractical to provide a recycle catalyst cooler on a unit of this scale merely for the purpose of being able to operate occasionally a t throughputs of 4 or 5 barrels per day. At this high feed rate the corresponding regenerator bed vapor velocity causes excessive overhead catalyst entrainment. In addition, at feed rates above 4 barrels per day the pressure drop in the reactor product lines becomes excessive, and the feed vaporizer and superheater heaters become overloaded. Throughputs a t the lower range are limited by having too low a velocity in the reactor riser or transfer line. Table I shows the range of capacities of the reactor and regenerator, and the factors limiting these capacities. INSTRUMENTATION
Instrumentation of a fluid type catalytic cracking unit is considerably more complex than that of most pilot plants because of the rather delicate pressure balance which has t o be maintained between the reactor and the regenerator vessels. The problem may be visualized by considering the junction of the regenerated
catalyst standpipe and the oil feed line on Figure 3 as a reference point. At this point the oil codd either flow up through the slide valve, the regenerated catalyst standpipe, and the regenerator bed or through the reactor riser and the reactor bed. The direction of flow will depend only upon the relative pressure drops across the two paths. Several methods of instrumentation are in use for solving this problem and the related catalyst flow and level control. Only the method found most satisfactory for this particular pilot plant is described here. It will be noted on Figure 3 that a differential pressure recorder controller (DPRC) mlintains a constant pressure relationship between the bottom of the reactor and regenerator beds by regulating the regenerator pressure. The flow will go in the desired direction-Le., up the reactor riser and into the reactor-
TABLE 1. CAPACITIES O F REACTOR AND REGENER.4TOR R A N G E O F CATALYST CIRCULATION RATES Reactor Oil Charge Rate, Bbl./Stream Day Low 1.5 Norm'al 2-3 High 4 Regenerator Carbon Burning Rate, Lb./Hr. Low 0.5
Normal High
1. 5 - 2 , 5 3
Catalyst Circulation Rate, Lb./Hr. Low 126 Normal 200-300 High 500
AXD
Reason Reactor riser velocity falls below 7 feet per sec., too low for transferring catalyst
......
Reactor product line pressure drop becomes excessive and charge heaters become overloaded
tween carbon content of spent and regenerated catalyst
......
Regenerator temperature rises above 1100" F. with all compensation windings cut o u t Catalys't bridges in standpipes
......
Determination of circulation rates carbon balance becomes inaccurate cause difference between carbon spent and regenerated catalyst is small
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I N D U S f R I A L A N D E N G I N E E R I N G CHEMISTRY
January 1948
if the pressure drop across the reactor riser is lower than the head of catalyst in the regenerated catalyst standpipe plus the pressure differential maintained by the differential pressure recorder controller instrument (this may be positive or negative, depending on operation and design). The head of catalyst above the bottom of the reactor and regenerator bed does not affect the balances at all. The setting of the differential pressure recorder controller will depend on the catalyst density in the regenerated catalyst standpipe, on the pressure drop across the reactor riser, and on the pressure drop across the regenerated catalyst slide valve. For a given catalyst, the reactor riser pressure drop depends on the catalyst and vapor flow rates. Using a riser velocity above 7 feet per second, the catalyst density is kept much lower in the riser than in the standpipe, and is in the range of 1 t o 3 pounds per cubic foot as compared t o standpipe densities of 30 t o 50 poynds per cubic foot. An equivalent pressure balance can be set up for the junction of the spent catalyst standpipe and the air inlet line. The one setting of the differential pressure recorder controller must assure proper flow directions in both transfer lines. Thus, it is possible that operating conditions which may appear sati'sfactory cannot be used in actual operations, because, regardless of the setting of the differential pressure recorder controller, by-passing may occur up either the reactor or regenerator standpipes. Control may be lost a t very low riser velocity when the density becomes excessive or a t very high velocity when friction causes excessive pressure drop in the riser. By-passing of feed or air t o the regenerator, the reactor, and stripper introduces a considerable explosion hazard and the following method is used to minimize the probability of bypassing. As indicated in Figure 3, a differential pressure recorder controller is placed across each slide valve. These instruments do not normally control the valves, but when the differential pressure of either valve approaches zero or becomes negative the instrument completes an electric circuit which opens a solenoid valve in the air supply t o both slide valves, thus closing them immediately. As soon as a normal differential pressure is re-established, the solenoid valve closes and the slide valve returns t o its normal setting. Three catalyst bed levels have t o be maintained in the pilot plant: reactor, stripper, and regenerator. With the total catalyst inventory fixed, only two of these need t o be controlled. AS shown in Figure 3, the reactor level is controlled with the re-
2.
*
PLAN VIEW OF SLIDE V A L . GATE
DIRECTION OF CATALYST FLOW Figure 5.
Catalyst Slide Valve
1. Guides for slide valve plug Motor valve stem
2.
3. 4.
Vessel standpipe Liner
5. 6.
7. 8.
Slide valve gate Slide valve plug Cooling fins Gasbleeds
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generated catalyst slide valve. Stripper level control is difficult without a slide valve in the transfer line from reactor t o stripper. Friction drop in this line being small, the stripper level will depend on the reactor level and on the difference between the reactor and stripper pressures. The catalyst capacity of the stripper is much lower than that of the reactor (6 us. 80 pounds), so that a small fluctuation in reactor level has a large effect on the stripper level unless it can be compensated for instantly by a change in either stripper pressure or catalyst flow rate. Ordinary mercury pot contmllers were found t o be much too slow, causing filter plugging and frequent loss of stripper level, which resulted in by-passing of stripper gases t o the reactor. The following control methods were found very satisfactory. The spent catalyst flow rate is maintained constant by a fixed setting of the spent catalyst slide valve. The differential pressure across the stripper bed is transmitted t o a Republic force balance type instrument which has a very low lag. This transmitter actuates the tube system of a pressure recorder controller which resets the stripper pressure control valve. This method of control is fast enough t o maintain the stripper catalyst level within a few inches. The heights of all catalyst beds are determined from two measurements: the differential pressure across a given vessel and the bed density. The latter is determined by measuring the differential pressure between two points within the bed separated by a known distance. When the vessel differentia1 pressures and cross-sectional areas are known, the catalyst holdups may be calculated. Direct measurements of catalyst flow rates by Venturi meters, slide valve pressure drop, or riser pressure drop are not very satisfactory. The most satisfactory method found for the pilot plant operation is based upon a carbon balance. From the difference in carbon contents between the spent and regenerated catalyst and a knowledge of the quantity and composition of the flue gas (which permits the calculation of the quantity of carbon burned per unit of time), the catalyst rate can be calculated. Feed and product flow rate measurement require much greater accuracy in the pilot plant than on a commercial plant if accurate weight balances are required. The oil feed drums are located on platform scales, Feed rates are set by varying the feed pump stroke, but the change in weight is used as the primary measure of feed rate. Injection steam (added t o the oil feed t o facilitate vaporization and increase the riser velocity) and stripping steam are also set by pump stroke. Water is drawn from graduated glass cylinders which serve as primary measures. The water is vaporized between the pump and the injection point. Liquid product is measured in accurately calibrated closed-top product tanka made of &inch pipe. Reactor and stripper gaseous products are passed through Monel wet-test meters. As in most pilot plants operating a t elevated temperatures, tempefature is controlled by electric heating elements wrapped around vessels and piping; this is in sharp contrast t o commercial units which operate with reactor temperature controlled by catalyst circulation rate, and regenerator temperature controlled by removing the excess heat of combustion in steam generating recycle coolers. In the pilot plant where radiation losses are high compared t o the heat content of the flowing streams, such control is not possible. The most critical temperatures on the pilot plant are controlled by an on-off type multipoint Celectray temperature controller, while the other temperatures are regulated with variable transformers which vary the voltages to the heating elements. More important temperatures are recorded on charts, while the others are read hourly on a multipoint indicator. MECHANICAL FEATURES
The mechanical aspects of the pilot plant discussed below are mostly concerned with features peculiar to fluidized bed units. It is believed that most of them will be applicable t o other processes in which the fluidized bed technique is used.
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HEAT COMPENSATION, INSULATION, AND HEATERS. The combination of high temperature level and high heat losses of the pilot plant makes it necessary to provide heat compensation to all lines and vessels. Furthermore, bring-up time is shortened by using the heat compensation elements t o hasten the attainment of reaction temperature on the unit. There is a total of approximately 123 kw. maximum in heating elements. These elements are made of Chromel wire of varying lengths and gages, depending on the rating required. The element wire is manually threaded into ceramic “fish-spine” type insulating beads. Control to each element is by means of an on-off switch or variable transformer. On the more important sections o€ the plant, such as the reactor and regenerator vessel and transfer lines, a combination of two or more elements of different ratings is used to give maximum flexibility. One or more of these elements will be in series with a variable transformer. The lines and vessels are insulated with several materials. (1) The pipe or metal is completely wrapped with asbestos cloth 0.125 inch thick; (2) the ceramic beaded elements are wound around the members; (3) the inner high temperature layer of Vermiculite insulation is added; (4) the low temperature insulation layer is added followed by a layer of light gage sheet metal for protection. Heaters are of the lead bath type with temperatures maintained by Chromel resistance elements placed around the outer circumference of each bath. LINESAND VALVES. The flow of abrasive catalyst makes it necessary to have streamlined piping with gradual bends and welded fittings. Lines as well as vessels a t 800’ to 1100O F. are of schedule 80, 4 to 6% chroinium4.5 molybdenum carbon steel, while those below this temperature range are mild carbon steel. Catalyst sampling lines and product lines which always have fine catalyst particles in them have cock valves. Where necessary, a continuous Bow of steam, air, or inert gas is passed through and/or around them in order to minimize their freezing with catalyst. A detail of a slide valve which is used in the control of catalyst circulation rate is shown in Figure 5.
A cylindrical slide valve plug having a pear-shaped gate moves back and forth past a circular opening in the bottom of the standpipe. The pointed end of the gate forms a 30” angle and permits close control of catalyst rates. The plug is actuated by a pneumatic diaphragm. The sleeve and plug are fitted to within a tolerance of 0.004 inch. Considerable erosion of the upper surface of the valve plug has been experienced a n d consequently hardened tool stev’ i s used to minimize this. The sleeve is constructed of 18-8 stainless steel. Two to 3 cubic feet per hour of bleed gas must always be flowing into each end of the valve between the sleeve and plug and into thc main catalyst stream. This prevents catalyst from collecting between these parts and freezing the valve action. All motor valves are of conventional plug and seat type and those subject to catalyst action, such as the product valves, require case-hardening of seats and plugs in order to decrease erosion. Thus, these parts will give 3 to 6 months’ service before requiring replacement. IKSTRUMEST LINE CATALYSTTRAP. All instrument lines taking off from zones where catalyst is present require a continuous positive back flow of gas through them in order to prevent catalyst from plugging them. ,For a 0.25-inch line, 2 to 4 cubic feet per hour of gas are adequate. I n the reactor it is important to keep the mole fraction of oil a t a maximum. Instrument line catalyst traps have been devised which not only reduce the required bleed gas to 1 cubic foot per hour but also give more positive assurance of lines not plugging as a result of momentary pressure surges. Figure 6 shows the details of one of these traps. The principal features are the flat Aloxite stone disk seated on the asbestos gasket and the valve spring compressed by the pipe plug. Bleed gas passes downward through the stone, thus keeping the stone and the 0.125-inch drilled hole below the stone free of any densely packed catalyst. GASSAMPLING ELEMENT.Obtaining catalyst-free gas samples Kas made possible by the device shown in Figure 6, which consists of a cyliqder of Aloxite filter stone cemented over a section
Vol. 40, No. 1
of perforated stainless steel tubing. The cement which withstood the high temperatures was a mixture of sodium silicate and talc. It was found advisable to enclose the stone filter in a metal cage to prevent the stone from dropping into the lines or vessels in case of breakage. Based upon the same principle as this small scale sampling tube, larger ones make it possible to procure gas samples from various heights and radial positions in commercial unit regenerators. These large sized sampling tubes are steam-jacketed in order to freeze the reaction a t the point of sampling. PRODUCT GAS CATALYST FILTERS.The product gases may entrain from 1 to 100 pounds of catalyst per 1000 cubic feet of gas, depending on operating conditions. In order to prevent plugging of product lines as well as high catalyst losses, a catalyst filter system is installed above all pilot plant vessels. I n Figure 7 are shown the details of the catalyst recovery system used above the regenerator vessel. The flue gas and catalyst first pass through a 3-inch cyclone where 50 t o 90% of the catalyst is knocked out of the gas. Then the gas with the remaining catalyst enters an Aloxite stone filter section where essentially all the remaining catalyst is separated from the gas. There are four thimble type stone filters individually supported by a stirrup and plate arrangement shown in the drawing. These stone thimbles have a permeability of 5 (cubic feet of gas flow per minute per square foot per 2 inches of water pressure drop) and are 36 inches long, 4 inches 0.d. and 3 inches i.d. At the top of each thimble is a machined surface upon which an asbestos gasket seats. The upper surface of the gasket seats into a machined groove of the flanged head. The four thimbles are located in one housing. Three thimbles are used t o filter gas while one is being purged free of catalyst by a reverse flow of purge gas, which may be steam, air, or nitrogen, depending on the nature of the gases being filtered. The flow through the thimbles is changed so that normally each thimble is purged once every 4 hours. During periods of extremely high loading, as is often experienced with fresh catalyst or a t very high bed velocities and heights, it is necessary to purge each filter as often as once every hour. OPER4TING SCHEDULE
Since it takes anywhere from 8 to 16 hours to bring up and line out the pilot plant, 8 hours a t lined out conditions before making a run and 8 hours during a run, the operation must be carried out 24 hours a day and 7 days a i.*eek. On 4 t o 12 and 12 t o 8 shifts there will be present one operator, his assistant, and one control tester. The pilot plant, therefore, requires 3 people per shift or
LNRECTION GAS FLOW
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Instrument Line Catalyst Trap
1. 2. 3. 4.
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Asbestos gasket 05-inch standard pipe thread 8. Drilled 0.125 inch 7.
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a total of 13 people (including shift breaker) on shift work. The special or long range analytical testing personnel required varies from 3 t o 6 per day, depending on the nature of the runs. Two project engineers plan runs in conjunction with the group leader and a general pilot plant supervisor. These engineers correlate data and work up operating reports and research reports based upon the data obtained from the pilot plant, and also keep abreast of general catalytic cracking developments. A calculating section works up yield structures, keeps track of incoming analytical
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Catalyst circulation is continued until the vessel temperatures reach 600" t o 800" F. At this time nitrogen is substituted for all air streams to the reactor and stripper. After the reactor and stripper are purged with nitrogen for 20 minutes, oil can be charged to the reactor, replacing the nitrogen to the reactor transfer line. Oil charge and injection steam rates are lined out through the vaporization section t o t h e oil drop-out line for some time prior to turning the oil into the reactor. With oil going to the reactor, heating element voltages and regeneration air are adjusted in order t o line out the unit. After 1 or 2 hours, control analyses can be started and operating changes made until the unit is in equilibrium a t thedesired conditions. SAMPLING, COLLECTION, AND WORKUP OF DATA
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Figure 7. 1. 2. 3. 4. 6.
G a s Sampling Device
*/winch stainless steel tubing
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0.75-inch pipe plug High tern erature cement Section oPperforated tubing 6. Aloxite filter stone
data, and tabulates data in suitable form in order t o facilitate reporting and correlation. This section also keeps a file of the original data sheets. The control tester runs those tests required to guide the operation of the unit with respect to catalyst rates and conversion level: flue gas analysis, carbon on the catalyst, and A.S.T.M. distillations of the total liquid product. Tests conducted on the day shift include gas gravities, mass spectrometer analysis of gaseous products, precision distillations of total liquid product, catalyst inspections, octane numbers, hydrocarbon type analysis, total sulfur, etc.
The data and samples required for a pilot plant run vary greatly, depending on what is being studied. As an example in a study of regeneration variables it may be unnecessary t o obtain a complete workup of hydrocarbon product properties. However, a certain minimum of data is collected in practically all runs and the discussion below is confined t o these data. CATALYST.Catalyst is generally collected from the regenerated and spent catalyst standpipes and from the reactor bottom. These samples are of considerable importance because the interpretation of practically all phases of the fluid bed catalytic cracking process depends on an accurate knowledge of catalyst properties. Catalyst has a tendency to stratify. Thus, two consecutive samples of spent catalyst may have slightly different carbon contents. Composite samples are therefore necessary for the study of a run. Variations in fines content are avoided by sampling after catalyst has been returned to the unit from the filter sections. If the hydrogen content of the catalyst is required, contact of the hot catalyst with air is avoided by sampling in a closed steel tube. It has been found that the catalyst carbon content is not affected by contact with air. Catalyst samples are taken every half hour before and during runs from the spent catalyst and regenerated catalyst standpipe for carbon analysis as a check on catalyst flow rate and regenerator performance. Information on catalyst properties normally required during a typical run is: Surface area (by low temperature nitrogen adsorption) Pore volume (by low temperature nitrogen adsorption) Particle size distribution (by Redimentation) Carbon content (every hnlf hour on spent and regenerated cstalyst, for control) Hydrogen content of coke Relative catalyst activity (dcterrnined in a fluidized fixed bed testing unit) Gas and coke producing factor (determined in a fluidized fixed bed testing unit)
OPERATING PROCEDURE
When the unit is started up, heating elements are turned on and all instrument lines and product lines are blown clear of catalyst. Air is charged to the regenerator and reactor through their respective transfer lines. Instrument bleeds are adjusted to the quantities to be employed during the run, and the addition of catalyst t o the regenerator is started as soon as the reactor and regenerator temperatures have reached 300" t o 500 O F, Catalyst is added from a storage hopper into the regeneration air stream. After the approximate quantity of catalyst required for the run has been added to the regenerator all pressures are put on instrument control and brought up t o the desired levels. The reactor level controller index is gradually moved to the required setting. This opens the regenerator slide valve and allows catalyst to move to the reactor until the desired level set by the index is reached. If more catalyst is needed, it may be added a t this time. Next, the stripper slide valve is partially opened, starting catalyst to flow from the stripper t o the regenerator. With the reactor level on control, continuous circulation is established in this manner,
REACTORAND STRIPPERGAS. Both gaseous products are sampled continuously during a run by displacing brine from a 10-gallon gas bomb. At the end of a run the gas is compressed slightly by connecting the bombs to a brine reservoir located on the second floor of the pilot plant. The gases are analyzed on a mass spectrometer for all components from hydrogen through the pentanes and amylenes; the quantity of hexanes plus is also obtained. REACTOR LIQUIDPRODUCT. An average of 13 gallons of composited total liquid product is retained from each run. Workup of this product varies with the nature of the study. The following properties are generally ascertained:
A.P.I. gravity A.S.T.M. distillation Podbielniak analysis Precision dist,illation into Ch fraction, depentanized gasoline, and cycle oil. Properties of these fractions are determined. Stripper liquid product is genSTRIPPERLIQUIDPRODUCT. erally recovered as an emulsion with condensed p i p p i n g steam
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ties of all oil and gas samples withdrawn are recorded, since they may be significant in the determinations of a weight balance. DATA OBTAINED IN PILOT PLANT
A Figure 8. A. B. C.
D. B.
F.
Regenerator Filter Housing
Gas catalyst mixture from regenerator T o cvclone Fro; cyclone Catalyst-free gas Rods and late shown turned 90° o u t of true position Catalyst Braw-off
stabilized by catalyst which leaks through the filter section'. The total emulsion from a run is collected and the oil separated by filtering through a'bed of stainless steel helixes or glass wool. Properties of the stripper oil are then determined after additional filtering and drying. FLUEGAS. Flue gas is analyzed by Orsat every half hour during a run t o check the catalyst rate. A slipstream from the flue gas line is taken directly into the buret of the Orsat located close by. TEMPERATURES. With temperature normally maintained by external heating rather than the heat content of the flowing streams, it is important that temperatures be obtained a t many places in the unit, because with improper compensation they may vary several hundred degrees over a very short distance. All ammeter readings (21) for compensating heaters are recorded hourly, Temperatures a t five different heights are recorded hourly for the reactor, stripper, and regenerator besides the outlet temperatures of the oil, air and steam heaters, transfer lines, standpipes, etc. PRESSURES. Pressure measurements are of considerable importance in continuous fluid bed processes. Besides vessel and heater inlet pressures, differential pressures are recorded to permit calculation of bed height, catalyst densities, and holdup in the reactor, stripper, and regenerator. Since a high degree of accuracy is required for differential pressure readings, water manometers are usually used. Where differentials are too high, Vizzene-B, having a gravity of 1.989 a t 79' F., is substituted for water. Slide valve differential pressures are measured with a water manometer; this serves as a rapid approximate check on the catalyst rate. FLOW RATES. Measurement of charge and products has been discussed under instrumentation. Readings are taken hourly. Xitrogen and air instrument bleeds, filter purge medium, and regeneration air are metered with rotameters operating a t 60 pounds per square inch. As many as eight bleeds are metered with a single rotameter. The manifolds are located adjacent to the instrument and the quantity to each bleed is set by the incremental rotameter readings. A water balance is made for all runs. Water, from the reactor and stripper phase separators is withdrawn periodically and the quantity is recorded. Quanti-
It is not within the scope of this paper to discuss in any detail the data obtained by the fluid bed catalytic cracking process. The list of studies which were made in this pilot plant is given above. I n order to present a general idea of the type of problem studied in the pilot plant, the following paragraphs discuss briefly the principal types of data obtained during studies of catalyst comparisons, comparison of operations a t 5 t o 10 pounds per square inch with those a t higher pressures, and catalyst regeneration. The comparison of the various types of cracking catalysts and catalysts of varying activity levels included Filtrol (treated naturally occurring silica-alumina catalyst), fresh and aged ground 3A (synthetic silica-alumina catalyst), and spraydried microspheroidal 3A catalysts. These studies were carried out so that process data such as product distributions and quality could be obtained. I n addition, there mere obtained engineering data such as catalyst activity decline rates, required cracking severities to give comparable conversion, carbon burning rates, attrition and entrainment rates, and relative flow characteristics. The comparison of operations at normally encountered pressures of 5 t o 10 pounds per square inch with those a t higher pressures from 25 t o 75 pounds per square inch were conducted mainly t o study the effects of increased pressure on catalyst activity decline rate, coke yields, and product quality. I n addition, great emphasis was placed upon obtaining relative cracking rates and carbon burning rates. The study of the variables affecting carbon burning rates involved not only pilot plan: runs made specifically t o study khis problem, but also the correlation of a great number of pilot plant runs which were conducted for the primary purpose of studying the reactor side. The regeneration studies involved the accurate analysis of flue gases, catalyst bed temperatures, holdup, and catalyst circulation rates as rvell as a knowledge of the amounts of carbon on the spent and regenerated catalyst streams. APPLICATION TO COMMERCIAL DESIGSAND OPERATIONS.The catalyst evaluations and the pressure and regeneration studies have been of considerable aid in designing and operating the commercial units of this company. The attrition rates, activity decline rates, and reaction rates determined on the pilot plant for the microspheroidal catalyst were big factors in the decision to employ this type of catalyst in the Houston commercial unit. The pressure studies which showed increased reaction rates and similar coke yields a t higher pressures were the basis for operating many units at their maximum pressure rating and for designing new units at pressures 10 t o 15 pounds per square inch higher than those designed prior t o the knowledge of what influence pressure would'have in the fluid bed process. The pilot plant regeneration studies made it possible in most cases to predict commercial carbon burning rates t o within *loyo. Thus, the magnitudes of the various operating changes required to achieve desired burning rates have been made available t o the operating staffs. I n addition to aiding in the operation of the regenerators, i t has also been possible to compare the relative performances of regenerators with different design features. Thus, for example, the relative efficiencies of grid type and spider type injection systems have been compared. ACKNOWLEDGMENT
The authors of this paper are greatly indebted to K. S. Sussan, J. T. Foulds, J. F. Crocoll, M. L. AndrB, H. G. Sandland, and other members of these laboratories for their contributions to the design and operation of this pilot plant. RECEIVED November 14.1947.
