93 1
Ind. Eng. Chem. Res. 1990,29,931-942
Development of Catalytic Cracking Technology. A Lesson in Chemical Reactor Design Amos A. Avidan* Mobil Research and Development Corporation, Paulsboro Research Labortory, Paulsboro, New Jersey 08066
Reuel Shinnar* Department of Chemical Engineering, The City College of New York, New York, New York 10031
Catalytic cracking converts heavy petroleum fractions into gasoline, distillates, and light olefins.
It revolutionized refining over 50 years ago and has been evolving continuously since. Catalytic cracking was commercialized in fixed, moving, and fluid-bed reactors (fluid catalytic cracking, FCC, has created fine powder fluidization). The design aspects of this important process, in three major reactor types, constitute a fascinating lesson in reaction engineering. We examined the properties of catalytic cracking, its chemistry, kinetics, and thermodynamics, and their influence on design. Of the hundreds of complex reactions, catalytic and thermal, it is the conversion of hydrocarbon t o coke and gas that dominates reactor and regenerator designs. The effects of hydrodynamics on cracking are examined through the concept of contact time distribution, and the discussion of the various reactor and regenerator types concludes with a summary of process control.
1. Introduction We are honored to participate in an issue dedicated to Hugh Hulburt. Hugh was a dear friend, and one of his great contributions to the profession was his editorship of the monthly, Industrial & Engineering Chemistry Process Design and Development. I&EC has published some of the most significant contributions to the understanding of catalytic cracking over the years, as can be judged from our references. Hugh had a deep understanding for the need to encourage publications that deal with the design process in a rigorous, academic way, and in this paper we analyze the development of catalytic cracking from a reaction engineering point of view. Catalytic cracking has been one of the key processes in petroleum refining for over 50 years. Modern fluidized solids technology has its origin in the development of fluid catalytic cracking (FCC) in the 19409. Catalytic cracking has evolved significantly in the past 50 years but has not reached the status of a mature technology. It was first commercialized in fixed-bed reactors using salt bath heat exchangers. The moving bed reactor, a widely used technology today, is derived from thermofor catalytic cracking (TCC) a moving bed catalytic cracker. However, today FCC has gained prominence (over 350 commercial units worldwide, with a capacity of over 1.6 X los m3/day). The history of the development and commercialization of catalytic cracking was reconstructed in detail in a recent review (Avidan et al., 1990). In this paper, we will analyze the development of catalytic cracking from the point of view of today’s reaction engineer. We want to show how the properties of catalytic cracking, and available technology, determined its development. 2. Chemistry of Catalytic Cracking The first, and one of the most important, part of a reactor design is the understanding of process chemistry. The complex processes that occur on a catalytic surface during the cracking process are poorly understood. However, for a reaction engineer, understanding the chemistry has a very specific meaning, and Hugh Hulburt was one of the first to clearly define it. What is needed is an understanding of the overall reactions, their thermodynamics, and the way overall reaction rates are affected by
* Correspondence can be addressed to either author. 0888-5885/90/2629-0931$02.50/0
Table I. Thermodynamics of Cracking Reactions (at 800 9) kcal/mol reaction AH AG 2 17 -9 3 12 -14 4 (n = 2) -12 14 5 18 -43 -15 -76 6
different process parameters. The level and detail of the required knowledge about process chemistry depend on the purpose. Developing and testing new catalysts requires a more detailed knowledge than is needed for the design of reactors. We will limit ourselves here with reactor design. Consider the cracking of a high molecular weight paraffin to a lower paraffin and an olefin: (1) CnH2n+2 CmH2m+2 + Cn-m&(n-m) +
for example,
-
(2) C20H42 ClOH22 + ClOH20 Cracking is endothermic (thermodynamic properties are given in Table I), a feature that has an important influence on reactor design. Naphthenes undergo similar reactions, while aromatic compounds can undergo side-chain splitting, yielding an olefin and an aromatic. Multiring aromatics normally do not crack but can undergo transalkylation and side-chain splitting reactions. The increase in the number of moles in cracking and equilibrium is favored by low pressure. However, conversion is constrained by kinetics and not by thermodynamics. There are several competing reactions to the type shown in eq 1. Paraffins and olefins formed can crack further by the same mechanism. Olefins can crack by a reaction such as C10H20 2C5H10 (3) Olefins can also polymerize n-C5H10 C5nH10n (4) and can also undergo hydrogen-transfer reactions such as C1oH2o CTH8 + C& + 2H2 (5) or CloH2o CYHs + 3CH4 (6)
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0 1990 American Chemical Society
932 Ind. Eng. Chem. Res.*Vol. 29, No. 6, 1990
t Yield. Wt%
Yield Vol%
Gibbs
5
0
€4
70
80
90
1W
€4
Con"err!on,VIllS
__
I
I
I
200
400 Temperature, "C
600
60 0
1 2
800
Figure 1. Effect of temperature on Gibbs free energy of cracking for a model compound.
Olefinic polymers and other high molecular weight compounds with low hydrogen content, plus low molecular weight gases such as CHI and H2, are formed. Some of these higher molecular weight compounds form "coke". Overall, one can write C,H,, cake f gas (7) It is this reaction that dominates the design of catalytic cracking reactors. One important question in reactor design is the temperature range that minimizes thermodynamic constraints on conversion or the range of the desirable reaction AG, which is significantly negative (preferably less than -5 kcal/mol). There is a large number of individual compounds involved in catalytic cracking, but some conclusions can be drawn from model compounds. Cracking reactions are endothermic, but as the AH is significantly larger than the AG, AG decreases with increasing cracking temperature. For one model compound (Figure l),AG is -5.0 at 300 "C. Generally, the lower the molecular weight of a compound, the higher is the temperature required to crack it. Thermodynamics thus restricts the lowest temperature for cracking for any catalyst to approximately 300-400 "C.
-
3. Kinetics of the Cracking Reactor Cracking is a complex process since there are thousands of individual compounds in the feed and in the products. Fundamental kinetic modeling, following all compounds, has not been performed as of yet and may be unnecessary. Many of the species present can be lumped into groups. The more lumps, the more difficult is the mathematical treatment and the more accurate is the description of the system that is needed (including feed analysis). Weekman (1969) has shown that, for reactor design purposes, a simple three-lump model is sufficient, as it contains the salient features of the system important for the designer: feed
kt
gasoline
(8)
\ / . coke
+ gas
Reactions 1 and 2 are pseudo second order (second order in mole fraction but not in partial pressure), whereas reaction 3 is pseudo zero order in mole fraction and therefore independent of conversion. Reaction rates depend on the catalyst used, operating conditions, and feed properties. All three reaction rates are affected by catalyst deactivation, which has a permanent component and a regenerable part (due to coking). One can write for atl three reactions kj = k 10. t cn-1 (9) where Iz,, is the catalyst activity for the regenerated catalyst, n varies from 0.4 to 0.6 for silica-alumina catalysts (n is lower for zeolites), and t , is the average residence time of catalyst in the hydrocarbon environment.
70 80 90 Canversion. VoI%
1W
Amorphous Catalyst Zeolhe Catalyst
Figure 2. Gasoline and coke yields vs conversion for zeolite (on silica-alumina support matrix) and silica-alumina (amorphous) catalysts (based on FCC kinetic model (Jacob et al., 1976)).
For coke formation, this results in the Voorhies (1945) equation: C, = At,"
(10)
The exponent n = 0.5 was related by Voorhies to a diffusion mechanism of coke precursors through a coke layer. In differential form, eq 10 is dC,/dt, = nAt,"-l
(11)
The average catalyst residence time, t , is simply related to the space velocity and the catalyst to oil (C/O) ratio by
Combining the principles of Voorhies (1945) and Blanding (1953), Weekman and Nace (1970) represented cracking as an apparent second-order reaction: dX/dt, = k,Q,(100 - X)'
(13)
where X is conversion, typically defined as 100 - volumetric yields of products boiling higher than gasoline (220 "C). Note that this "conversion" is really a selectivity. Cracking is a high conversion process in the sense that almost every molecule in the feed undergoes some change, but conversion as used here is typically 70-80 vol %. Q, is a catalyst decay function. Two forms of Q, can represent cracking data: a first-order decay (exponential) and an nth power decay similar to eq 9. In addition, it is assumed that the cracking activity declines in proportion to the coking activity. Furthermore, Weekman and Nace assumed that the decay function for gasoline cracking (k2) is of the same shape as eq 13 with the same but with a different rate constant. Many more complex mechanisms have been proposed, but the mechanism of eq 8 is sufficient to illustrate basic principles. One of the most useful lumped models published, and the one we use here too, is the one published by Jacob et al. (1976). A ratio of coke yield to normalized conversion (or "crackability") has been used for many years to analyze cracking data: k , = coke yield/crackability = C(lO0 - X ) / X (14) The typical gasoline yield over zeolite and Si/Al catalysts in gas oil cracking is shown in Figure 2A as a function of conversion. The gasoline yield increases with conversion up to a maximum and then decreases as reaction 2 in eq 8 predominates, and gasoline cracks to lighter products and to coke (coke yield is shown in Figure 2B). 4. Cracking Catalyst Deactivation by Coke
Presently known cracking catalysts deactivate quickly by coke deposition. Some coke is formed by polymeriza-
tion and condensation of primary cracking products. Coke
Ind. Eng. Chem. Res., Vol. 29, No. 6 , 1990 933
I_\ \ i
Ratio HIC Aiomic
530
6 : HIC Ratio in Combusted or Stripped Coke
A. HIC Ratio as Function of 1.Hexane Conversion (Adapted from Wolciechowski and Coma, 1986)
520 Temp. ‘C Required tor Constant Conversion 510
HIC Atomic Ratio
500
0 1 0 z a 3 0 4 0 ~ 1-Hexane Convenlon pi%)
Figure 3.
H/C ratio in coke.
10000
1000
Cracking Catalyst Activity
100
5
0
Stripping or Combustion Time
10 15 Coke Content on Catalyst, %
20
Figure 5. Effect of coke content on activity of Pt reforming catalyst.
r
2
I -
Coke, % on Catalyst
1
Amorphous Silica-Alumina Silica-Alumina 1 0
1 Coke Content on Catalyst, %
2
Figure 4. Cracking catalyst activity vs coke content.
is not a clearly defined compound but a complex mixture of hydrocarbons. Coke also contains sulfur and nitrogen compounds, which are the source of SO, and NO, emissions in the regenerator flue gas. While it is often assumed that coke is mostly carbon, in the form of combined aromatic rings, other compounds are also present. A considerable fraction of the coke is easily strippable by steam, or other gases, at typical reaction temperatures (540 OC). At higher temperatures, gasification by steam may also occur. The history of catalytic cracking has been dominated by the competition between cracking and coking. Some fraction of the coke deposited on spent cracking catalyst comes directly from unconverted feed molecules. These can include complex structures such as asphaltenes. A large fraction of the Conradson carbon residue (CCR) in the feed is incorporated into coke. The typical H/C atomic ratio of coke for one model compound (1-hexene) is shown in Figure 3A for a common cracking catalyst, a Y zeolite. The initial coke formed has a H/C ratio of 2, a ratio close to the H/C ratio in the feed and the major product-gasoline and olefinic hydrocarbons. As conversion increases, the coke becomes more aromatic. The high H/C %eke" is the first to be removed by stripping or by combustion (Figure 3B). The more “graphitic”, ‘condensed”, or “residual” carbon is removed last, at a slower rate. High-efficiency combustors can reduce residual carbon on regenerated catalysts down to less than 0.05 w t %. The cracking activity decreases with the coke content of the catalyst. A typical plot for pure zeolite and an amorphous silica-alumina catalyst is given in Figure 4. During burnoff, the coke is redistributed, and residual coke, while reducing activity, has a lower impact as compared to fresh coke. For comparison, the activity of a reforming catalyst as a function of coke content is shown in Figure 5. The effect of coke on activity is much smaller in reforming, and the catalyst can have a much higher coke content before most activity is lost (20% versus 1% in catalytic cracking).
0 0
10
20 30 Time on Streams, Sec.
40
50
Figure 6. Coke deposition on zeolite and silica-alumina catalysts as a function of time.
Coking occurs rapidly, and the coking rate is a function of catalyst properties, feed composition, temperature, and pressure. For a given temperature, pressure, and feed composition, the coke level is simply a function of time. A typical coke-time curve is shown in Figure 6 for silicaalumina and zeolite catalyst. Comparison of Figure 6 to Figure 4 shows that zeolite catalyst loses over 90% of its initial activity after 1s, whereas silica-alumina loses this amount of activity in over 1 min. Coking deactivates a cracking catalyst within a much shorter time period than in almost any other catalytic process. The catalyst therefore requires frequent regeneration, and it is this frequent regeneration that dominates the design. There is another aspect of catalyst regeneration that dominates design consideration. As approximately 5 wt 9% of the feed is converted to coke, it is important that the heat evolved in combustion be recovered. The other mechanism of catalyst deactivation is “permanent” deactivation by dealumination and metal poisoning. This component of deactivation cannot be recovered in the cracking reactor-regenerator system. Dealumination is caused by steaming in the regenerator, which is proportional to steam partial pressure, regeneration temperatures (in an Arrhenius form), and catalyst activity to a power larger than one. Metal poisoning by sodium, nickel, and vanadium is more important as heavier feeds (such as residuum) are processed. Metals buildup on the catalyst increases coke and gas yields and accelerates catalyst deactivation. Catalyst makeup requirements in catalytic cracking are determined by the need to maintain a desired activity level to compensate for these deactivation mechanisms. 5. Thermal Cracking
Another important competing reaction is thermal cracking. Thermal cracking was first used for the production of gasoline from fuel oil 80 years ago and is still used today in resid cracking, visbreaking, and coking op-
934 Ind. Eng. Chem. Res., Vol. 29, No. 6, 1990
erations in petroleum refineries. Thermal cracking is less selective than catalytic cracking as more methane is formed and the ratio between reactions 1 and 7 is lower. Thermal cracking is undesirable, and the reactor designer must minimize it. The rate of thermal cracking strongly depends on temperature and feed. Below 450 "C the rate of thermal cracking is low for most hydrocarbons, while above 600 "C the rate is fast for all hydrocarbons except methane. Some compounds tend to undergo thermal cracking, forming coke quite fast above 475 "C. Such compounds are lumped together as CCR. The race of thermal cracking increases with molecular weight and therefore with the boiling point. On the other hand, thermal cracking may play a positive role in the catalytic cracking of heavy petroleum fractions. Very large molecules cannot be catalytically cracked in a zeolite catalyst, but fragments and free radicals formed by thermal cracking can be acted upon. However, for lighter fractions, thermal reactions are detrimental to product quality and yield. To summarize, the main features of catalytic cracking chemistry are as follows: (1) The reactions are consecutive, and the desired products such as gasoline continue to crack to lighter products and coke. (2) Coke and low molecular weight gases such as H2 and methane are formed as undesirable parallel products. (3) Thermal cracking occurs at high temperatures, and the feed cannot be preheated above 425 "C. Some cracking of gasoline into lighter gases is not necessarily detrimental, as the olefins and isobutane produced can be alkylated. The first major use of catalytic cracking in World War I1 was to produce aviation gasoline, which was composed of FCC naphtha, alkylate, and tetraethyllead. Increasingly, light olefins today may be valued more than gasoline for petrochemicals production. The development of cracking technology has been dominated by the need to minimize undesirable side reactions such as thermal cracking and coking while maximizing yields of desirable products. One way to achieve that is the continuous search for better catalysts. Over the years, development of new catalysts led to substantial improvements in the relative rates of cracking to thermal reactions, increasing gasoline yield. Synthetic silica-alumina catalysts followed natural clays, and in the early 1960s Mobil introduced zeolites that increased reaction rates and selectivity (up to 50% higher gasoline yield). These developments had a significant effect on unit design. A second way is to develop better reactor designs, and it is this way that is our main concern here. To understand the choices facing the designer, we will first discuss the impact of operating conditions such as pressure, temperature, and residence time. 6. Effect of Pressure
Pressure has little effect on the overall thermodynamic constraints within the range of low pressures used in catalytic cracking. Lower hydrocarbon partial pressure favors cracking over coking and is therefore desirable. The minimum total pressure is determined by the cost of the gas plant and is at present above 300 kPa. However, hydrocarbon partial pressure can be reduced by steam injection (typically 1-5 wt % of the feed). Other means to reduce partial pressure include the recycle of some light hydrocarbon gases, but the cost should be balanced against the yield benefits. A typical plot of the effect of pressure on k, (coke yield divided by crackability) is shown in Figure 7.
1.8
1.6
1.4
1.2
1
100
150
200
250
300
Gas-Oil Feed Paltlal Pressure, kPa
Figure 7. Effect of pressure on coke yield (based on FCC kinetic model (Jacob et al., 1976)). 1.5
I
500
480
520
540
Temperature, 'C
Figure 8. Effect of temperature on coke yield (based on FCC kinetic model (Jacob et al., 1976)). 100
Mid-Continent Gas Oil Feed
Rate Constant lSec
1 ° T 10.3
1.0
1.1
1.2
1.3
'4
17(IPK x 1 0 3 )
Figure 9. Thermal cracking data (Chen and Lucki, 1986).
7. Effect of Temperature Cracking temperature is one of the main variables that allow the operator to change yields and product composition. For the designer, three effects are important. (1) The activation energy of coking is lower than that of cracking. This effect can be seen in Figure 8 where coke yield divided by crackability is plotted as a function of temperature. There is an advantage to operate at higher temperatures, as this reduces the amount of coke forlned for a desired conversion. (2) The limit of the permissible temperature (aside from metallurgical problems) is due to thermal reactions. Thermal cracking by noncatalytic free-radical reactions has a higher activation energy than catalytic cracking and therefore predominates at high temperature. Thermal cracking rate for Mid-Continent gas oil feed is given in Figure 9 (Chen and Lucki, 1986). (3) Higher temperatures require shorter residence time (higher space velocity) to prevent overcracking, and there are practical limits to how low they can be reduced to. 8. Residence Times a n d Contact Times
To reduce thermal cracking, the reaction engineer has another tool-the relative adjustment of contact versus
Ind. Eng. Chem. Res., Vol. 29, No. 6, 1990 935 residence times. When two parallel kinetic reactions are promoted by the same catalyst, selectivity can be improved by proper choice of a catalyst. However, if one of the two reactions is thermal, one can also minimize the actual residence time, while keeping catalytic conversion constant. Catalytic conversion is determined at a given temperature and pressure by the following. (1) Space Velocity. The space velocity is the hourly mass flow rate of the feed divided by the total weight of catalyst in the reactor (WHSV). Decreasing the space velocity while keeping the residence time constant increases the ratio of catalytic cracking to thermal cracking. However, there is a limit as to how much catalyst can be fitted into a given volume. In addition, selectivity suffers because of increased solids backmixing. (2) Activity of the Catalyst. The ratio of catalytic to thermal cracking can be increased by increasing the activity of the catalyst. If the ratio of thermal to catalytic cracking is fixed, then a given catalyst activity per unit volume results in a maximum permissible temperature and a maximum permissible residence time. Higher temperatures reduce the ratio of catalytic coking rate to catalytic cracking rate and also improve some product properties, such as octane. (3) Uniform Contact Time Distribution. The contact time distribution (Shinnar et al., 1973; Shinnar, 1986) is a dimensionless measure of how conversion and product distribution in the reactor deviate from plug flow. Uniform contact time improves selectivity. The concept of a residence time distribution has had a strong impact on reaction engineering over the past 30 years. But it only applies rigorously to homogeneous reactors, as it assumes that all molecules stay in the reactor the same amount of time regardless of their state. In heterogeneous systems, molecules are absorbed at a surface, and adsorption and desorption are selective: the real sojourn times of molecules can be different for different species. Furthermore, it is not residence time in the reador that counts but the time in contact with a catalytic surface. For a first-order reaction, one can define a contact time distribution rigorously. If we assume a first-order reaction network with rate matrix Ri, then the outlet composition in a plug flow reactor is Cj = Ajo CAjnCXnr (15)
+
where Aj,, are constants determined by the feed composition, A,, are the eigenvalues of the rate matrix, and 7 is a time constant related to space velocity. For any isothermal catalytic reactor with arbitrary flows, one can show in a similar fashion to a homogeneous reactor that outlet composition is Cj = Ajo CAjn@(Ant) (16) The funciton $ ( A n t ) is the Laplace transform of the function W). The contact time distribution is only correct for the specific pseudo-first-order reaction matrix for which it was measured. It is a quantitative measure of how the flow deviates from plug flow. As molecules are absorbed on a catalyst surface (or inside zeolite crystals), the actual sojourn time of a molecule in the reactor can be much larger than the residence time in the gas phase. Depending on absorption, the difference can be an order of magnitude. There is virtually no data in the literature to account for this effect, which is very important in a fixed or fixed-fluid bed reactor. On the other hand, in an entrained-bed reactor such as an FCC riser, the sojourn time of any molecule is limited to the residence time of the catalyst particles. (Typically, the average solids residence time in the FCC reactor is 3 times
+
Relative Burning Rate
1.1
1.0
1.2
1.3
irr (IPKx 103) Figure 10. Coke burning rate as a function of temperature (based on Weisz and Goodwin (1963) and Prater et al. (1983)). 35
'"h 25
ResIdence The, Minutes
600
650
700
750
Temperature,' C
Figure 11. Effect of temperature on coke burning rate.
the average gas residence time, or 5-25 s.) Therefore, if thermal reactions affect molecules absorbed on the Catalyst surface, then there will be a significant difference between an entrained bed and a stationary catalyst type of reactor (such as many laboratory-scale FCC reactors (Sapre and Leib, 1990)). The flow pattern can deviate considerably from plug flow, if catalyst densities vary across the cross section or if the catalyst phase has significant backmixing due to clustering. \k(t)an be strongly different from plug flow, which can lead to yield penalties. A uniform \k(t)also requires that there would be no diffusional resistances inside the catalyst particle. This limits the particle size. 9. Kinetics of Coke Combustion Catalytic cracking yields large amounts of coke (4-8 w t % of feed). Hydrogen-rich compounds either volatilize at the temperatures of a typical regenerator (650-800 "C) or crack to combustible volatiles and coke. The time required for catalyst regeneration is dominated by the rate of 'slow coke" combustion. The activation energy of the combustion is approximately 35 kcal/mol. A typical curve of reaction rate versus temperature is given in Figure 10. The time required for regenerating a catalyst with 1% coke (with 1%oxygen breakthrough) is given in Figure 11 as a function of temperature. This time scale is important as it fixes catalyst inventory and regenerator size. Reducing the size of the regenerator and the catalyst inventory are important for two reasons: FCC regenerators are large and their cost is a significant fraction of total unit cost. Furthermore, minimizing catalyst inventory has significant advantages not only in reducing operating costa but also in allowing faster adjustment of the inventory to changing needs of feed or product composition. FCC units operate in "heat balance", and hot catalyst from the regenerator supplies part of the heat of cracking. The FCC unit heat balance depends on catalyst activity, feed properties, feed preheat, and reactor temperature. The heat balance also depends on the desired C02/C0 ratio in the regenerator flue gas. The products of the
936 Ind. Eng. Chem. Res., Vol. 29, No. 6, 1990 Flue Gas
Product
Catalyst Maximal Particle Size, mm
4
7egeneration
Cracking Regenerated
500
550
600 650 Regeneration Temp., "C
700
750
Figure 12. Maximal particle size to avoid intraparticle temperature excursion. Air
primary combustions are CO and C02 and HzO, and the C02/C0 ratio is a function of temperature. CO reacts with O2to form C02in a thermal free-radical reaction, but this reaction is slowed by the presence of solids. If the flue gas contains excess air, it will ignite as soon as the solids are removed. Until recently, the design and operation of regenerators was dominated by this feature. Nowadays, additives containing precious metals, such as platinum, allow complete combustion to C02. The combustion of CO to C02 can also be achieved by increasing combustion temperatures. As the activation energy of homogeneous combustion of CO to C02 is high (approximately 70 kcal/mol), temperatures of above 700 "C will result in complete CO combustion, given sufficient air, but material and catalyst stability constraints reduce the maximum temperature below that obtainable from the heat balance. Catalysts deactivate at high temperatures by sintering and by steaming. Today's FCC catalysts can be safely regenerated at temperatures of up to 850 "C without excessive damage to them by sintering, but deactivation by steam requires significantly lower temperatures. The designer therefore faces a compromise between minimizing steam deactivation by lowering temperature and reducing the size of the regenerator by increasing it. Another key to reduce deactivation of steam is to use two-stage regeneration: the first stage, where H2-richcoke is burned, is operated at low temperature; the second stage is operated at a high temperature. The required residence times for combustion given in Figure 11were calculated for an isothermal vessel. FCC regenerators are not always isothermal, and one potential source for temperature nonuniformity is mass-transfer limitations inside catalyst particles. The impact of mass transfer has to be minimized so that the temperature inside the catalyst particles will not exceed gas-phase temperature. Diffusion control in combustion is a result of fast reaction rates. Diffusion times vary with the square of the diameter of a catalyst particle, whereas reaction rates are a function of temperature. Figure 12 shows the relation between the maximum particle size and temperature. Particle sizes need to be below 200 Fm so that the regenerator can operate above 650 OC while avoiding high intraparticle temperatures. The minimal particle size in fixed bed reactors is 1 mm whereas in the moving bed it is about 3 mm. Only fluid beds operate with particles below 200 Fm. In processes that have much lower coke deposition per unit of feed, there are other ways to protect the catalyst. A common way is to use dilluted combustion air. This reduces the adiabatic temperature rise of the air-carbon reaction but is impractical in FCC due to the large volumes of air involved.
Feed
Figure 13. Schematic diagram of continuous activity reaction system (FCC,TCC).
10. Development of the Cracking Reactor Having outlined the chemistry of cracking reactions and the kinetic constraints, let us now examine how different designs dealt with the challenges. We want to show why modern FCC technology won over the other designs and examine the chemical reaction engineering lessions that can be learned for similar design problems. Three different types of reactors have been developed for catalytic cracking. The fixed-bed reactor enjoyed prominence from 1936 to the early 1940s but is no longer used today. Only two reactor types are still in use, and only one, the fluidized bed catalytic cracker (FCC), is used in new designs. Few operating moving bed reactors (TCCs) are left today, and they are being gradually replaced by fluid-bed crackers. But all three designs have had a lasting impact on chemical reactor design technology. To understood the underlying reasons for reactor choice, let us examine the criteria resulting from the kinetic constraints discussed in the previous sections. (1) Catalytic Cracking Using a Catalyst That Deactivates Quickly Due to Coking. In a fixed-bed reactor, each new element of feed meets a catalyst with lower activity than the previous element of feed. As conversion depends on the product of catalyst activity and space time, this is similar to a reactor with constant catalyst activity but nonuniform contact time. This lowers conversion and the yield of intermediates in a system of consecutive reactions. One way to overcome changing the catalyst activity is to move the catalyst continuously through separate reaction and regeneration zones (Figure 13). The average reactivity of the catalyst bed is kept constant by continuously removing coked catalysts from the bed and adding regenerated catalyst to it. Each catalyst particle still experiences a reaction-regeneration cycle, but the feed sees constant average catalyst properties. Two approaches to realize this concept of moving catalyst between reactor and regenerator were developed simultaneously around 1940. Both used the conceptual scheme illustrated in Figure 13, and both are still being used for many other applications today. One used a moving bed with large catalyst particles; the other used fluidized particle technology, which can either be a dense fluidized bed or an entrained bed. (2) Effective Heat Transfer and Utilization. The cracking reactor must be able to transfer heat to the reaction at temperatures of 500-550 OC. Cracking generates large amounts of coke, and the heat of combustion must be removed in a controllable way while utilizing the heat
Ind. Emg. Chem. Res., Vol. 29, No. 6, 1990 937 of combustion efficiency. In principle, one can use most of the heat of combustion for steam generation and power recovery and supply the heat of cracking separately. But complete thermal independence between reactor and regenerator not only complexes the design but also lowers the overall thermal efficiency. It is more efficient to utilize a large part of the heat of combustion in a direct way. This has lead to the "heat-balanced" concept in FCC where the catalyst particles carry sensitive heat to vaporize, preheat, and supply the heat of cracking. Part of the heat of combustion is recovered from the flue gas, and part is lost. The FCC unit is heat integrated on the reactor side as well, and the feed in preheated by heat exchange with products. Other heat-exchange means, such as fired feed preheater, and an external cooler to remove heat from the regenerator are also used to provide flexibility and controllability. The regenerator is not always in balance with the heat requirements of the reactor. It is therefore desirable to design the system so that heat can be both added and removed independently. The regenerator itself should be operated as closely to isothermal conditions as possible. This requires efficient heat removal. If combustion rates are fast due to high temperatures, the catalyst particles have to be small to avoid excess temperatures due to diffusional limitation inside the particles. Heat can be removed in two ways: by heat exchange inside the combustion zone, using either water, molten salt, or molten metals, or by using catalyst flow to remove heat from the regenerator to the cracking reactor. Fluidized bed technology has a large advantage in heat removal. Fluidized beds have intense mixing of the solids and therefore high heat-transfer coefficients and close to isothermal operation. In a tubular reactor, the temperature nonuniformity in the catalyst bed is a function of diameter, local flow rate, and local reaction rates. This puts strong limitations on the maximum allowable local reaction rates. The requirements of efficient heat removal and heat addition are often the dominating parameters in the choice of the reactor. The heat-transfer requirement, together with the requirement of catalyst reactivation, dominated the development of hardware and catalyst technologies. The designs that were developed in the past 50 years form the basis for many other present reactor designs. (3) Operating at High Temperatures. Operation at higher temperatures only became desirable for high-activity catalysts that were not available at the time of different designs developed. Short residence times are possible in fixed beds, but short residence times at high temperatures require very rapid heating of the feed. If the heating of the feed is done by hot catalyst particles, then the temperature of the particles is significantly higher than the averge cracking temperature. Mixing must be rapid to avoid thermal cracking. This strongly limits the particle size, as the heating time of 3-mm particles is about 2 s. (4) Operating with Changing Feed and Changing Product Requirements. Cracking feedstocks can vary widely, and the cracker must be able to handle a wide variety of feedstocks. Product demands vary seasonally. Modern fluidized bed crackers are able to meet these demands by changing catalyst type or activity, reaction temperature, and space velocity. Changing feed properties and conversion also affects the heat balance, as coke make and heat requirements of the reaction change. In the next sections we will discuss how different designs dealt with these issues. 11. Fixed-Bed Reactor The first commercial catalytic cracking reactor was a
PRODUCT
I STOCK CHARGE STOCK PREHEATER
FLOW SHEET, HOUNDRY PROCESS
Figure 14. Schematic diagram of the fixed-bed catalytic cracking unit. A~xAIR
AND OIL VAPOR OUTLET I
CATALYST FILLER PIPES
AI I/
TOP TUBE SHEET-
77m
REGENERATIO
.5m CATALYSTTUBE SHEET
I l l
VAPOR HEADER
Tm
MAIN TUBE SHEET
lm
SALT OUTLETS
VAPOR TUBES-
(55) LEGEND AIR OR VAPOR PATHS
a
SALT PATHS
0CATALYSTSPACE
L
L SOIL A LVAPOR T i N INLET L E T AIR AND
NOTE DRAWING IS NOT TO SCALE DISTANCES ARE APPROXIMATE ONLY A FEW TUBES ARE SHOWN FOR THE PURPOSE OF FOLLOWING THE OIL FLOW SMALL ARROWS INDICATE PATH OF AIR AND OIL VAPORS THRU THE CASE
Figure 15. Diagramatic sketch of commercial Houdry case.
fixed bed (Figure 14). It was jointly developed by Eugene Houdry, Sun Oil, and Socony Vacuum, the forerunner of Mobil. The first 2000-BPD unit was started up in April 1936 in the Paulsboro, NJ, refinery. Its reactors contained acid-treated natural clay catalyst. Three reactors were operated in parallel, and they were switched from cracking to regeneration every few minutes. The heat transmission between cracking and regeneration was achieved by a molten salt heat exchanger. The molten salt was flowing through the reactor in a series of tubes that were equipped with fins to increase heat transfer (Figures 15 and 16). The molten salt extracted the heat from the reactor operating in the combustion mode and transferred it to the reactor operating in the cracking mode. The fixed-bed reactor was an ingenious solution to meet the requirements of a cracking reactor with the technology available then. It represented a major milestone in process technology, including such highlights as the first wide-
938 Ind. Eng. Chem. Res., Vol. 29, No. 6, 1990 AIR
TUBE SIZES 1. REGENERATIONTUBES
-
OUTSIDETUBE INSIDETUBE
-
-
Scm. 3cm.
2. DISTRIBUTORTUBES
-
OUTSIDETUBE INSIDETUBE
-
-
6cm. 3cm.
3. COOLINGTUBES
-
OUTSIDETUBE INSIDETUBE
-
6cm. 3cm.
4. COOLINGTUBES
-
3.3cm.
-
TO REACTOR
/ STEAM
HEATCONDUCTING FN I S-
I
TUBE
TlON
I BLOWER THERMOCOUPLE
1
ARRANGEMENT
AND OIL VAPORS ENTER CATALYST BE13
t
u
Figure 18. Schematic diagram of the TCC air lift process.
(AIR AND OIL VAPORS LEAVE CATALYST BED THROUGH COLLECTOR TUBE ORIFICES) NOTE DRAWING IS NOT TO SCALE
Figure 16. Cross-section sketch of typical tube unit layout for commercial Houdry case. Instantaneous Yield 50
Gasoline Yield,
r-11 1 I\\ I 35t \\ I
45
40
Amorphous Catalyst
30
60
I
I
I
70
80
90
\\
ELEVATOR VESSEL
I
100
Figure 17. Integrated gasoline yield in a fixed-bed reactor vs instantaneous yield achievable in a constant activity reactor (based on FCC kinetic model (Jacob et al., 1976)).
spread use of catalysis in petroleum refining, air regeneration of coked catalysts, and the implementation of automatic process control. however, considering the demands of the cracking reactor given in section 10, the fixed-bed reactor had shortcomings. (1)Fractions of feed saw an increasingly deactivated catalyst. The first fraction was overcracked, whereas the last fraction had low conversion. This is similar to a reactor with a broad residence time distribution. Conversion is integrated over different histories, which reduces intermediate yields in consecutive reaction schemes. This is illustrated in Figure 17, where the integrated yield is compared to the instantaneous yield. If catalyst deactivation is slow, one can compensate for it by raising the temperature. This is impractical in catalytic cracking, as deactivation is too fast and the temperature is already high. The large thermal holdup of the salt bath would not allow such a control scheme. This would be possible if the time scale of the deactivation was measured in weeks instead of minutes. (2) Temperature distribution was not uniform during combustion, thus limiting the maximal regeneration temperature.
The other requirements listed in section 10, namely operating at high temperatures and short residence times, and the ability to handle a large variety of feeds, were not as important in the late 1930s. The two competing cracking technologies were developed to improve on item 1, the ability to handle a fast-changing catalyst. Today's salt-bath tubular reactors are descendants of this technology, and in practically all of them, the catalyst is contained in tubes. Most of them operate with long times between regeneration (2 weeks to a year). There is no need to utilize the heat evolved in the regeneration. Only one major process still operates with very frequent switching between reaction and regeneration-catalytic dehydrogenation of butane and propane (Catofin and Cadadiene). This reactor is a direct descendant of the Houdry reactor, but also has moving bed and fluid-bed competitors. The moving bed process, UOP's Oleflex, has been introduced recently. It is based on the moving bed CCR reformer, which in turn is based on TCC principles. Fluid-bed technology has been practiced in the Soviet Union and Romania. 12. Moving Bed Reactor (TCC)
A schematic representation of a TCC unit is given in Figure 18. The moving bed reactor uses large catalyst particles (3-5-mm spheres). The catalyst moves by gravity from the reactor section to the regenerator section, and the regenerated hot catalyst is lifted by an airlift to a solids-air separator from which the catalyst is fed by gravity to the reactor. A small side stream is fed to an elutriator to remove fines generated by attrition. The TCC had many advantages over the fixed-bed reactor and allowed much larger units. It solved one of the major problems of the Houdry reactor, namely, constantly changing catalyst activity. The mechanical problems of such a system were formidable but were solved by elegant design. This technology has found important new uses such as in UOP's continuous catalytic reformer (CCR) and Oleflex processes. While the TCC had large advantages over the fixed-bed reactor, it still fell short in meeting the requirements given in section 10. The large particle size limited the regenerator temperature to below 650 "C, requiring a large regenerator and large catalyst holdup (Prater et al., 1983).
Ind. Eng. Chem. Res., Vol. 29, No.6, 1990 939 1.o
Per
lh Height
1.o
Typical Gas Velocities for Fine Powders (mh):
-
Bubbling 0.1 Turbulent 0.5 Fast 3 Riser 10
-
0.5
Non-Dimentional
0.1
.
-
Solid Volume Fraction
Turbulent
0.01
0.001
Bubbling I \
0 0
0.2
I
0.4 Solid Volume Fraction
___
I
0.6
n.ow1 0.001
0.8
Figure 19. Axial density profiles in fluid beds.
(Modern TCC catalysts allow regenerator temperatures of up to 675 "C, in properly designed kilns.) Furthermore, it was difficult to transfer the heat of combustion to the reaction zone. Due to the large particle size, the difference in temperature between the reactor and the regenerator had to be minimized. The required difference is a function of catalyst to oil (C/O) ratio. In early (bucket elevator) TCC units, this ratio was limited to 1.5, the air-lift TCC reached 4.5, and the later designs reached a value of 7. C/O ratios of 9 or more are possible in FCC units. This meant that the TCC regenerator had to be cooled by steam generation, and the feed had to be vaporized. Later TCC designs would accept liquid feed but still require a high preheat of a t least 400 "C. TCC reactor temperature was typically 500 "C and could be raised up to 525 "C. On the positive side, the moving bed with a vapor feed has very uniform contact time distribution (provided that reaction rates are low enough so that particle diffusion is not controlling). A moving bed design falls short a t meeting the requirements of high temperatures, short residence times, and changing feed properties. The large particle size limits the rates of heat and mass transfer. The large catalyst holdup increases the response time of the system to catalyst changes, which limits the ability of the cracker to handle variations in feedstocks and product demands. These shortcomings were not as important 50 years ago as they are today. There were no active catalysts to allow high temperature and short residence times. Another disadvantage of TCC technology was not related to reaction engineering but to mechanical design problems. The cost of most process equipment per unit feed decreases with increasing capacity. However, for each unit there is a critical size at which the cost per unit starts to increase with size. For a TCC this happens a t about 30000 BPD. FCC's have strong cost advantages for larger sizes. The average size of an FCC unit has increased over the years to about 40000 BPD, and units as large as 130OOO BPD have been constructed.
13. Fluid-Bed Reactor (FCC) An alternative way to realize the flow scheme shown in Figure 13 is to use fluid-bed technology. FCC catalyst consists of spherical particles, with a particle size varying from 1to 130 Fm. When a fluid is passed through a bed of solids, the pressure drop increases, and when it exceeds bed weight, the solids start to mix and behave like a fluid. At higher velocities, bubbles form and then undergo continuous breakup and coalescence until the bed reaches the turbulent fluidization regime. As gas velocity is raised further, the boundary between bed top and freeboard disappears and the entrainment rate increases. The fluid bed is maintained by feeding back the entrained particles.
0.01
0.1 10 10 Superficial Gas Velocity ( mis)
100
Figure 20. Modern high-efficiency FCCU with riser reactor and riser regenerator-fluidization regimes (from Squires et al. (1985)). FLUE GAS VENT
REGENERATED
MODEL I SCHEMATIC
Figure 21. Schematic diagram of the upflow model I FCCU (1942).
4
PRODUCTS TO SEPARATION TOWER
DISTRIBUTION GRID
ONTROL
STEAM PREHEAT FURNACE YCLE
RESIDUE
Figure 22. Schematic diagram of the Model IV FCCU (1952).
Typical axial density profiles are shown as a function of gas velocity in Figure 19. It should be noted that there is a transition between a dense turbulent bed to a more dilute entrained bed. There is a gradual increase in top density and a decrease in bottom density. Most academic research has concentrated on bubbling beds, whereas most commercial interest has been in "bubbleless" fluidization (Squires et al., 1985). The high-velocity, bubbleless regimes are used for cracking and regeneration, as shown for a modern high-efficiency FCC unit (Figure 20). Fluid beds have certain advantages over moving beds. (a) They can handle smaller catalyst particles. (b) Solid circulation rates can be adjusted over a wider range. (c) Fluids and solids can be mixed quickly, and rapid solids movement greatly improves heat transfer. An early FCC design is shown in Figure 21, which gives a schematic representation of the first Easo Model I upflow
940 Ind. Eng. Chem. Res., Vol. 29, No. 6: 1990
design (1942), and the 1952 vintage Model IV is shown in Figure 22. Model I was an “upflow” system, while Model IV was a “pressure-balanced downflow” system. Model I, which was the first commercial fluid-bed system using a fine powder, had many advanced features that were dropped in later models but eventually were reintroduced. It had upflow cracking and regeneration, quick separation of cracked products from spent catalyst, and catalyst coolers and feed preheat (”non-heat-danced” operation). Model I and most later FCC designs (except for Model IV) share some basic features: Catalyst flow rates can be adjusted over a wide range, allowing for great flexibility in opertion. In this paper we are not concerned with a detailed discussion of the design changes that led to todays different FCC units, but rather in the way the demands outlined in section 10 effected the design. Let us start with the regenerator 14. FCC Regenerators
Early FCC regenerators operated at tempertures below 650 “C. At these temperatures, there was little CO combustion in the gas phase. The flue gas ignited once the solids were removed from the gas in the cyclone due to the presence of CO and oxygen, and the absence of solids that quench free radicals. This afterburning could lead to excessively high temperatures in the cyclones, and oxygen conversion needs to be high, leaving less than 0.2% oxygen in the flue gas. Such high conversions are difficult to achieve in low-velocity fluid beds without a catalytic combustion promoter (if low carbon level on regenerated catalyst is desired). There is significant bypassing in a large bubbling bed, especially in shallow regenerator beds that were designed to minimize pressure drop and air blower costs. The practice of operating with low fines concentrations (because of low efficiency cyclones) caused large bubbles and aggravated the problem. Fluid-bed regenerator efficiency can be improved by operating at higher velocities such that there is a significant catalyst concentration in the freeboard. A typical FCC regenerator operates in turbulent fluidization at a gas velocity of 1 m/s, such that about 20% of the catalyst is in the freeboard. In fact, whole bed circulates through the cyclones about every 5 min. This solids holdup in the freeboard allows relatively high conversions in these shallow, inefficient regenerators. However, afterburning in the cyclones can still be a problem. The temperature rise due to afterburning is used to control the regenerator. To achieve higher CO conversion, one needs to provide combustion to COz either by raising the temperature (above 700 “C) or by utilizing a catalytic combustion promoter. Small amounts of Pt allow complete CO conversion at temperatures of 650 “C. The activation energy of CO combustion with a catalyst is only 20 kcal/mol (Prater et al., 1983). Even early FCC regenerators benefited from nearly isothermal conditions. Hot spots, which can easily develop in fixed-bed or moving bed reactors (there is no lateral mixing in TCC), are eliminated in fluidized beds. For the same reason, fluid-bed chemical reactors (such as acrylonitrile) can be operated in the explosive regime, without ever exploding. Still, fluid bed reactors are not completely isothermal. Lateral and vertical temperature differences can reach 50 “C, and larger temperature fluctuations can occur over very short time periods. But the impact is much less than that of excess temperatures inside a large catalyst particle. FCC regenerators underwent many changes over the years. One particularly successful design is the “high-
efficiency”regenerator shown in Figure 20. The turbulent fluid bed at the bottom of the combustor is turned into an entrained fast fluid bed by reducing the vessel diameter. The solids are separated from the gas in an inertial separator in the upper combustor. Part of the regenerated catalyst is returned to the base of the riser reactor, and part is recirculated to the lower combustor. This design resulted in a smaller vessel and better gas-solids contact inherent in high-velocity fluidization regimes. The highvelocity regenerator is more forgiving toward initial distribution problenis and allows reduction in unit inventory. It also allows complete CO combustion with little excess air and little or no promoter. Another new regenerator type uses two-stage combustion. The fast-burning hydrogen-rich compounds are combusted first at lower temperature in the first regenerator, and the slow-burning residual coke is combusted at a high temperature in the second regenerator. This concept has been commercialized in some resid cracking FCCU’s. The goal of this design is to reduce the exposure of the catalyst to steam generated in the combustion, as the hydrogen content of the coke and therefore the steam content in the flue gas are lower in the high-temperature zone. The two-stage design reduces catalyst makeup and increases the flexibility of the unit. 15. FCC Reactor The Model I FCC reactor was an upflow pipe or riser, with an expanded diameter section designed to provide more residence time. The upflow mode is also used today but at much shorter residence times. The designers of this early fluid-bed system soon discovered that fine powders do not seem to obey Stokes’ law. Even at velocities several times higher the single-particle terminal velocity (which is approximately 0.1 m/s for FCC catalyst), the bed is not entrained immediately. The particles behave as clusters with an effectively higher terminal velocity due to interparticle forces. This realization led to the development of Model I1 and subsequent “downflow”fluid-bed reactors. The expanded reactor section was elongated, a dense fluid bed was formed, and the solids were withdrawn from the bottoms, rather the top of the bed. The higher bed density in these downflow system allowed long contact times and lowered catalyst losses which were high in Model I. Improvements in catalysts, brought about by introduction of synthetic silica-alumina to replace natural clays, followed by high alumina catalysts, caused more of the cracking reaction to occur in the riser leading to the reactor. Mobil’s invention of zeolite-containing cracking catalysts in the early 1960s has eliminated the need for the dense bed altogether. Existing FCC reactors were modified to all riser cracking, usually by extending the riser through the reactor vessel. The “reactor”vessel, which is still called reactor in many installations, now houses the riser cyclones and, in some designs, the catalyst stripper (which uses steam to remove some adsorbed hydrocarbon from spent catalyst). The modern FCC reactor is different from conventional catalytic fluid-bed reactors. The catalyst is introduced through a standpipe, and oil feed is introduced into the bottom of the riser section. This allows rapid mixing between the vaporizing feed and hot regenerated catalyst. The bottom of the riser is a complex environment. There are strong axial and radial temperatures and solid concentration profiles. The desired goals for the mixing zone are rapid mixing of solids and oil and minimizing catalyst holdup in the zone. A high catalyst holdup indicates that vaporization is slow and that, therefore, the oil residence time is larger than indicated by average gas velocity.
Ind. Eng. Chem. Res., Vol. 29, No. 6, 1990 941 b) Time-Temperature Regimes for Fluid-Solids Systems
2000 r
Major Control Loops (Other Loops Omitted for Clarity) Reactor
Regenerator
UJ
QC Reaction
I--
To Main Fractionator
A-P +
(P+ I D)
Set
Temp.,lmE ‘F
Art 500
I
I
I
I
.
1.0 10 100 Residence Time, Sec.
0.1
1
SolldS
Figure 23. QC reaction system (from Gartside (1989)).
Reactor design in modern FCC units still falls short of the goals outlined in section 10. The achievable residence time is still larger than desired to achieve the optimum ratio of chemical to catalytic cracking at high temperatures. The limitations are in the gas-solids separation as well as in the bottom mixing zone. Recent advances in quick gas-solids separation have moved the main constraint into the bottom mixing zone. Vaporization and cracking require milliseconds for an active catalyst, but it is difficult to mix large amounts of catalyst (over 60 tons/min in a large unit) with cold oil feed within milliseconds. There is therefore a strong variation in the temperature-time activity history that different elements of the feed undergo. In that sense, the effective contact time distribution is highly nonuniform. One can measure the impact of this nonuniformity by comparing the overall results with that of a small pilot plant. The differences are important enough to merit improvement, but they are not large enough to outweigh the other advantges of this design. Faster mixing and shorter residence times are still an area where there is still a place for novel designs. One interesting development is the “quick contact” (QC) system described in Figure 23. This sytem can contact feed and catalyst, react them, and separate and quench them in less than half a second. (Current systems need more than 1 s.) However, no commercial plants are yet in operation. Understandingand modeling highly nonisothermal mixing systems in which residence times and mixing times are of comparable magnitude is still a challenge for the reaction engineer. There is another interesting lesson for the reaction engineer. The conditions at the bottom zone of a modern FCC riser are far beyond our capabilities to realize detailed scaleup by predictive modeling. This was less severe a problem in earlier designs as catalysts were less active and temperature differences were much smaller. We have to realize that sometimes successful designs involve risks. Such risks can be minimized but not avoided. 16. Design for Controllability One important consideration in the choice and design of a reactor is the ability to control it (Shinnar, 1981; Lee and Weekman 1976). In a reactor that has the potential for multiple steady states, it is also important that the dynamic control be able to maintain unconditional asymptotic stability. In addition, the cracker has to handle a wide variety of feedstocks and to change the product distributions and product specification, upon short notice. FCC’s ability to accept a wide range of feedstocks makes it the kingpin in controlling the whole refinery. Changing FCC operating conditions allows changes in the ratio gasoline to heating oil, when demands change from winter
Generation
Air Furnace Oil Fuel
Figure 24. One FCC control scheme with steam generation in the regenerator (Shinnar, 1981).
to summer. It also allows trade-offs between gasoline yield and octane. The fluid-bed unit has a greater degree of flexibility than both the fixed-bed and moving bed reactors. The degree of flexibility is a strong function of design and varies strongly in different designs. The critical items which allow FCC to operate over a wide range of conditions are as follows. (a) Catalyst. Continuous (or daily) catalyst addition and withdrawal of equilibrium catalyst compensate for permanent deactivation. The catalyst change in FCC is much more rapid than in TCC due to the lower inventory. The whole inventory in FCC is typically replaced every 50-100 days, and response to a new catalyst or additive is much quicker. (b) Reactor Temperature. Temperture is an independent variable that adjusts steady-state catalyst activity and allows changing product slate and properties (such as gasoline octane). (c) Catalyst Circulation Rate. Catalyst circulation rate can be adjusted over a wide range in most FCC designs. In a truly adiabatic FCC, in which there is no feed preheat and no catalyst cooler, the circulation rate for a given reaction temperature is mainly a function of regeneration temperature and therefore strongly constrained. If a preheat furnace is added to the unit, the catalyst circulation rate can be adjusted over a much wider range. Catalyst coolers are also used to adjust FCC heat balance. Catalyst type also affects circulation rate. A coke-selective, ultrastable catalyst would not change the coke yield in a heat-balanced unit but would lower the carbon on spent catalyst. The C/O ratio will hence increase. (d) C 0 2 / C 0 Ratio. Another important factor impacting the heat balance is the C02/C0 ratio in the regenerator. If no combustion promoter is used, this ratio is solely a function of regenerator temperature. Pt promoter allows operation at full CO combustion, at any operating temperature in most units. If one desires to operate at partial combustion, a CO boiler is needed to complete the combustion of the flue gas. Use of a promoter allows fine tuning of this ratio and thereby control of the heat balance. These and other control items are combined into an overall control scheme, such as the one shown in Figure 24 (Shinnar, 1981). The two main variables that would affect the conversion are reactor riser top temperature, which controls the catalyst circulation rate, and feed preheat temperature. One important control variable, available in the first Model I FCC unit and recently res-
942 Ind. Eng. Chem. Res., Vol. 29,No. 6,1990
urrected, is heat removal from the regenerator. Steam is generated in coils or an external catalyst cooler. This arrangement is especially useful when processing heavy feeds, such as resids with high CCR levels. Modern FCC design meets most of the criteria outlined in section 10 well and definitely meets most of those criteria better than any of the other designs. There are still two areas in which FCC designs fall short. (1)The lower limit of residence time is still higher than desirable, mainly due to limitation of the bottom section and the gas catalyst separation. (2) The contact time distribution is not uniform. Furthermore, the temperature activity history of different fluid elements can vary considerably. These shortcomings still leave room for future imaginative design. But any design is a compromise, and the present designs are better compromises than former technologies. We hope that the previous discussion is useful for the student in reaction engineering in teaching him that reactor design has to take into account a large variety of factors and always involves compromises.
Acknowledgment The figures in this paper are intended to demonstrate trends rather than show exact data. We are grateful to Raymond Hu for his assistance with Figures 1 , 2 , 7 , 8 , 11, 12, and 16. Lee Breckenridge supplied us with Figure 5. Figures 4 and 6 are drawn after several authors, including Voorhies; orders of magnitude was taken from Mobil researchers Farber, Payne, and Sailor (1965). Most other figures are adapted from Avidan et al. (1990).
Nomenclature A = constant in Voorhies equation (0.2-0.9 in ref) Aj, = constant determined by feed C = coke yield, wt % of feed C, = carbon deposited on catalyst, wt 70 (carbon on spent carbon on regenerated catalyst) CCR = Conradson carbon residue, or continuous catalyst regeneration (when applied to reforming) Cj= outlet reactor composition (eq 16) C/O = catalyst to oil weight ratio FCC = fluid catalytic cracking H/C = hydrogen to carbon atomic ratio in coke I&EC = Industrial & Engineering Chemistry k, = coke yield divided by crackability, X kj = catalyst activity for each reaction shown in eq 8 k,, kj, = catalyst activity for regenerated catalyst
n = Voorhies exponent, about 0.5 in Voorhies’ work Ri = rate matrix for first-order reaction network (eq 15) t , = catalyst residence time, min TCC = thermofor catalytic cracking WHSV = weight hourly space velocity X = conversion (=lo0- vol % yield of products heavier than gasoline, typically above 220 OC) X = Crackability, (=X/(1- X)) AG = change in Gibbs free energy AH = change in enthalpy @ = catalyst deactivation function A, = eigenvalues of the rate matrix Ri T = time constant related to WHSV in eq 16 ‘;k(t)= contact time distribution *(t) = Laplace transform of Jr(t)
Literature Cited Avidan, A. A.; Edwards, M.; Owen H. Innovative Improvements Highlight FCC’s Past and Future. Oil Gas J. 1990,88 (2), 33-58. Blanding, F. H. Znd. Eng. Chem. 1953,45 (6), 1186-1197. Chen, N. Y.; Lucki, S. J. Ind. Eng. Chem. Process Des. Deu. 1986, 25 (3), 814-820.
Gartaide, R. J. QC-A New Reaction System In Fluidization VI; Grace, J. R., Schemilt, L. W., Bergougnou, M. A., Eds.; Engineering Foundation: New York, 1989. Jacob, S. M.; Gross, B.; Voltz, S. E.; Weekman, V. W., Jr., AZChE J . 1976,22, 701-713. Lee, W.; Weekman, V. W., Jr. AIChE J. 1976,22,27. Prater, C. D.; Wei, J.; Weekman, V. W., Jr.; Gross, B. In Advances in Chemical Engineering; Wei, J., Ed.; Academic Press: New York, 1983; Vol. 12, pp 1-60. Sapre, A. V.; Leib, T. M. Translation of Laboratory FCC Catalyst Characterization Tests t o Riser Reactors; Advances in FCC Technology; American Chemical Society: Washington, DC, Aug 1990.
Shinnar, R. Chem. Eng. Commun. 1981,9,73-99. Shinnar, R. Use of Residence and Contact Time Distributions in Reactor Design. In Chemical Reactor and Reaction Engineering; Carbury, J., Varma, A,, Eds.; Marcel Dekker: New York, 1986. Shinnar, R.; Glasser, D.; Katz, S . Ind. Eng. Chem. Fundam. 1973, 12, 165. Squires, A. M.; Kwauk, M.; Avidan, A. A. Science 1985,230 (Dec 20), 4732, 1329-1337.
Voorhies, A., Jr. Znd. Eng. Chem. 1945,37 (4), 318-322. Weekman, V. W., Jr. Ind. Eng. Chem. Process Des. Deu. 1969,B (31, 385-391.
Weekman, V. W., Jr.; Nace, D. M. AIChE J. 1970,16 (3), 397-404. Weisz, P. B.; Goodwin, R. D. J. Catal. 1963,2 (5), 397-403. Wojciechowski, B. W.; Corma, A. Catalytic Cracking;Marcel Dekker: New York, 1986. Received for review October 6, 1989 Revised manuscript received January 18, 1990 Accepted February 22, 1990
