Jet Fuel Selectivity in Hydrocracking Bruce E. Stangeland' and James R. KittrelP Chevron Research Go., Richmond, Calif. 94802
The selectivity for producing jet fuel in single-stage, extinction recycle hydrocracking of a raw gas oil over a supported metal catalyst i s described by a series reaction model which represents selectivity changes resulting from kinetic temperature changes with a nonfouled catalyst as well as from temperature adjustments to offset loss of activity with catalyst fouling. Nonisothermal data of both types are used in the model specification and parameter estimation. A nonselective fouling model i s used wherein all individual reactions are equally inhibited during catalyst deactivation. The pilot plant data exhibit large declines in jet fuel on unfouled catalyst with increasing per pass conversion and smaller declines with increasing space velocity. Yields are relatively insensitive to increase in catalyst temperature during catalyst fouling.
P r o v i d i n g an adequate kinetic model for a reacting system of industrial importance is a worthwhile but elusive goal. A model of an industrial process is recognized as valuable, not only for process optimization, but also for new plant design. However, the complexity of industrial reactions often inhibits experimental modeling efforts on other than pure compounds. I n a hydrocracking reactor, for example, a mised phase system of hydrogen and hydrocarbon is cracked over a supported metal catalvst to lower-molecular-weight products. -4typical hydrocracking catalyst is a complex system of several metals on an acidic support. The hydrocarbon feed to be cracked can be represented as a spectrum of pure hydrocarboiis of molecular weights ranging from 100-500 and of varying proportions of paraffins, aromatics, and naphthenes; a variety of catalyst poisons nil1 also be present in the feed. Each of these feed compounds reacts n i t h varying speed and selectivity to jointly form a spectrum of products of lower molecular weight. This product spectrum is separated into specified boiling ranges to provide individual products of commercial importance such as jet fuel, gasoline, and butane. The stoichiometry of the overall reaction is unknown and represents a summation of many individual reactions. Furthermore, catalyst poisoiis deactivate the catalyst during operation, so the hydrocracking catalyst temperature must be increased continuously to maintain catalyst activity. With such complexity, neither a completely descriptive mechanism nor a completely accurate mathematical representation of the process can be achieved with reasonable expenditure of time and money. Kevertheless, the importance of selectivity in reacting svstems demands adequate kinetic modeling of relative product yields. In the present paper, we present both a modeling method and the results of an application of the method to the kinetics of jet fuel production in single-stage hydrocracking. The model represents the major trends in the data quite accurately nhile being no more sophisticated than the data warrant. Indeed, one major criterion in the model development must be the maiiitenance of minimum model complexity (parsimonious parameterization) while still providing an adequate data representation (Kittrell, 1970). Correspondence should be addressed to 691 Cragmont Ave., Berkeley, Calif. 94708. Present address, Department of Chemical Engineering, University of hlassarhusetts, Amherst, Mass. 01002.
The model to be generated must describe two apparently different modes of dependence of jet yield on catalyst temperature as typified in Figure 1. Firpt, the dependelice of jet yield on interacting variables such as conversion, pressure, and temperature must be described for a nonfouled catalyst. In particular, the model must be capable of describing the fresh catalyst yields a t 2000, 1600, and 1300 p i g for a variety of conversions and space velocities. Such changes in operating conditions on fresh catalyst are referred to here as kinetic changes. The 2000-psig yield data eshibit an ob erved temperature O F . The 957, conslope of (-2.67 f 0.56) X 10-I wt fidence region (assuming appropriate error properties of jet yield) thus indicates that this slope is significantly different from zero. Second, the model must simultaneously describe the 1300-pig data, which were obtained by operating aii Isomas catalyst a t constant conversion and space velocit'y as the catalyst temperature was adjusted to offset t'he effect of catalyst fouling. The observed temperature dependence of these (lata is (-2.12 + 2.4) x wt 7 , / O F and is sigiiificantly different from that of the 2000-psig data. The model developed here (or any other model for this process) must dmcribe both sets of data and differentiate between yield decline 1%ith kinetic changes and decline with fouling. Possible kinetic models of the deactivation reactions have been discussed (Szepe and Levenspiel, 1968). The effect of acbivity decay on reaction selectivity has also received theoretical treatment by Froment and Bischoff (1961 , 1962)
m-
-5 :u 0 -
m
-
o
0 =
Lrn L - .
45
40 720
F r e s h catalyst I
4
730
740
I
I
I
I
760 770 780 A v e r a g e C a t a l y s t T e m p e r a t u r e , "F
750
I
I
7W
8W
;; Figure 1. Dependence of yields on temperature for fresh and fouling catalysts
0 2000
psi
(> 1600 psi
0 1300 psi Ind. Eng. Chem. Process Des. Develop., Vol. 1 1 , No. 1 , 1972
15
and by Weekman (1969). I n this paper we describe the effect of deactivation on the overall reaction selectivity for supported metal catalysts during changes in temperature. For example, if the kinetics of the primary reaction can be written
k
=
-S,ln
(1 - X )
and then one could interpret catalyst deactivation as a continual decrease of a with time on stream. Hence, the catalyst temperature must be increased to offset the declining a! and to maintain the conversion, X, constant a t constant space velocity, S,. T o specify the selectivity of the overall reaction network, several secondary reactions generally must be considered; each reaction has its own individual preexponential factor, ai, which is decreasing with time. Changes in reaction selectivity with fouling could be ascribed to varying rates of decay of t'he a I ,for example, through varying rates of decay of hydrogenation and acid activities in hydrocracking. We propose a simple kinetic model that describes all of the yield data on fresh catalyst as well as fouled catalyst without reference to different rates of decline of the a t . Experimental
subsequent distillation to provide a narrow boiling range jet fuel. Substantial amounts of lower-molecular-weight gasoline and gases were also formed in the reaction. The maximum jet fuel yield t,hat can be obtained from hydrocracking depends on process conditions, the boiling range of the jet fuel, a i d the boiling range and origin of the hydrocracking feedstock. An arbitrary, narrow boiling range jet fuel is used throughout this paper as a reference jet fuel cut and is so designated in the figures. The feed for the tests was a broad boiling range, raw gas oil blend sufficiently characterized for the present study by its relative molecular weight and density; the ratios of jet fuel density and molecular weight to t'hose of the feed were, respectively, 0.882 and 0.544. The standard cat'alyst particle size for the studies was 8-14 mesh; however, one run was made a t 1300 p i g with 42-60 mesh catalyst to evaluate potential intraparticle diffusioiial limitations. This smaller mesh size had no effect on either the activity or the selectivity of the catalyst; the data are among those used in the subsequeiit anal Iluring one period the hydrogen rate was doubled. S o effect' on activity or selectivity was observed other than that expected from the concomitant changes in ammonia and hydrogen partial pressure. Hence, bot,h intraparticle and interparticle diffusional effects were judged to be negligible. Model Equations
All esperimental studies were conducted in a small laboratory pilot plant having a nominal fresh feed rate of about one-half gallon per day and closely simulating a commercial hydrocracking unit in that it is equipped with gas recycle, extinction recycle of unconverted feed, continuous distillation of product, and quantitative recovery of light products. Makeup hydrogen was added on demand to replace that consumed in the hydrocracking reaction, thus maintaining system pressure a t the desired level. The reactor wits maintained a t an isothermal condition well within a 2'F spread over the bed. The hydrocracking catalyst was a commercial Isomax catalyst and was brought on stream under conditions similar to those employed commercially. The schematic process flow diagram is shown in Figure 2 . Fresh feed was coinhined with liquid recycle and then pumped to the reactor inlet, mixed with hydrogen-rich recycle gas plus makeup, and passed through the reactor. The product wvas then flashed in a high-pressure separator a t system pre::Ywre. Recycle hydrogen was returned to the reactor inlet; electrolytic hydrogen was added as required to maint'ain system pressure. Liquid product from the high-pressure separator was depressured and fractionated in highly efficient distillation columns, giving separations comparable to T B P distillations. All unconverted splitter bottoms were continuously recycled to the reactor (Figure 2). The entire gas and liquid recycle system was designed for minimum holdup. The distillation column temperatures were set to provide jet fuel as the heaviest boiling liquid product-that is, in the viciriity of 500°F. The initial boiling point of the jet fuel was set b y
In this model the fixed bed catalytic reaction initially is considered describable b y power law kinetics. Adsorption terms will be included in the rate expressions only as required by the data. If the products are art,ificially divided into the products of commercial importance, various reactions that might participate are
F
ka
+U j L
T o simplify these equations, nonesseiit'ial products were eliminated from Equations 3. For example, the reaction with rate constant k l must produce lower molecular weight products in addition to jet fuel J ; these are not shown. I n this paper, only kiiletics of formation of jet fuel will be considered; hence, let us tentatively entertain the following simple set' of rate equations:
(5)
If the initial moles of feed and jet are F o and Jo, then the molar yields of unconverted feed and jet fuel are
(7)
Dislillatlon
Retlor Ff
column
Total Feed, Separator Unconverled Feed, 11-XI Ft
Figure 2. Schematic drawing of pilot plant configuration 16 Ind.
Eng. Chem. Process Des. Develop., Vol. 1 1 , No. 1 , 1972
Equation 6 assumes that feed disappears by first-order kinetics. The first exponential in Equatioii 7 represents cracking of jet in the fresh feed, the second formation of synthetic jet, and the third cracking of synthetic jet. These equations iu molar units are now converted to the more commonly used units of weight fraction yield to fresh feed, volume fraction of jet initially in the fresh feed, and conversion.
Since these equations refer t o the entire reaction mixture passing through the reactor, F o and J o must refer to the inlet composition of the tot,al feed. The total feed, F 1 , and fresh feed, F,, are related through the definition of conversion, X (Figure 2) :
X
= F,/I."i
(8)
The only jet, in the total feed is that entering with the fresh feed, which is Xo (volume fraction) of the fresh feed. Thus, the volume of jet elitering the reactor and based on total feed is
Jo, = XoF,
=
(1 - X , X ) F l
Fo
=
=
XoXFtpj,'A1[j
(1 - Xo.Y)Flp,..!llfp
where CJs represents the synthetic jet yield aiid differs from CJ by a coiistaiit
(10)
These two equations must be converted to a molar basis for use iii Equations 6 aiid 7 :
Jo
Equation 15 relates the overall reaction rate constant, k , to process variables. Substituting Equation 15 into Equation 16 gives
(9)
and tlie rest of the total feed boiling above the recycle cut point represeiits Fo:
Fo,
Model 1: k2 = ka = ks = 0. I n this case, k = kl and light gases and gasoline are produced only as accompanying jet fuel in the kl reaction. Here, Equation 14 simplifies to:
(11) (12)
Tlie filial objective is a i i expression for C J , thc n-eight of jet produced per weight of fresh feed: (13) Equatioiis 11-13 can he substituted into Equation 7 to obtaiii the desired espressioii for C J . If, in additioii, the residence time for both feed and jet, iii t,lic react'or c:in be approximat,ed by t = l/St,,then
The plot of k vs. reciprocal temperature shown in the upper portion of Figure 3 indicates no obvious deviation from liiiearity. Hence, there is 110 reason to doubt the validit'y of Equation 15. However, given t'his relationship for k , Equation 16 simplifies to Equation 17. Kote that this equation demaiids t,liat jet yields be independelit of temperature for both kinetic changes and fouliiig changes in catalyst temperature. Coiisequently, the temperature-dependent yield changes of Figure 1 require rejection of this single-parameter model. Model 2 : k2 = 0. T h e iiuinber of independelit parameters can be increased to t\vo by setting k 2 = 0 in Equation 14; this would appear more reasonable than adding a parameter to describe a reaction order or an absorptioii t'erm, since Equation 1.5 was not inadequate in Model 1. I n hlodel 2, Equation 14 simplifies t'o
PF
This apl)roxiniatioii works surprisingly nell, coilsidering the complcs nature of trickle bed reactors. The vapor/liquid ratio i i i c w w s with temperature as well as wit'h distance donii t,lie bed. ;Ifter t'he volume of t'he recycle stream is similarly converted to molar units, an expression for 6 from Equations 6 and 12 become,
E l ) ,jet cracking increases more than does jet production. If EP < E l , there mould be an increase of jet yield with increasing space velocity a t constant conversion. In Figure 7 the presence of jet fuel boiling range material in the hydrocarbon feed influences jet yield, but not on a one-to-one basis. When this jet fuel overlap increases from 6-30%, the jet yield does not improve b y 24% but only by about 7% a t low conversion. Selectivity losses with Catalyst Age
Improvement of fresh catalyst activity and maintenance of catalyst selectivity during catalysts aging have been goals sought b y cat'alyst'formulators for many years. These characteristics of commercial catalysts are due to a variety of factors. For esample, loss of jet yield with hydrocracking catalyst age has been ascribed to predominant loss of either hydrogenation or acidic activity; the resulting imbalance in activity of these two functions from the carefully specified values for the fresh catalyst, then, causes deleterious effects on yields. Ind. Eng. Chem. Process Des. Develop., Vol. 1 1 , No. 1 , 1972
19
I n Figure 8 the predicted yield decline with catalyst age is shown for two additional cases: k 2 is three times greater than the value used for Figure 3, and El is one-half as great as the value for Figure 3. Xpparently, the increase in k 2 serves primarily to decrease the absolute level of jet yield; a catalyst possessing high selec,tivity for cracking lower-molecularweight compounds could produce this effect. The decrease in E l , b y contrast, primarily influences selectivity changes during fouling. Such a decrease in E l could result either from diffusional limitations on the large feed molecules (which are absent for the smaller jet molecules) or from a refractory feed inherently exhibiting a lower El relative to E2. In any event, the selective site deactivation concept need not be invoked. Acknowledgment
The authors thank Chevron Research Co. for permission to publish this viork. Nomenclature
CJ
= total yield to feed of a narrow boiling range jet fuel,
CJa
=
@,,
=
E
=
E i= F
=
FI = Fo = I.'t = FoL,=
pi= G
=
J
=
wt fraction yield t'o feed of synthetic jet fuel produced by cracking higher-molecular-weight feed, wt fraction predicted value of total yield t'o feed of narrow boiling range jet fuel from X o d e l 3 , wt fraction activation energy of a reaction with rate constant k , kcal!g-mol activation energy of a reaction with rate constant ki, kcal/g-mol symbolic designation of gas oil feed or molar flow rate of this component, mol/hr flow rate of fresh feed to the recycle loop, cc/hr flow rate of unconverted total feed to reactor, mol/hr flow rate of total feed to the reactor, cc/hr flow rate of unconverted total feed to the reactor, cc/hr predicted value of the yield of unconverted feed from Model 3, wt fraction symbolic designation of gasoline product symbolic designat'ion of jet fuel product or mole flow rate of this component, mol/hr
flow rate of jet fuel in the total feed to the reactor, mol/hr J o ~ flow rate of jet fuel in the total feed to the reactor, cc/hr k rate constant for disappearance of gas oil feed, hr-1 ki rate coilstant for cracking in any individual reaction i , hr-* L symbolic designation of butane and lower-molecularweight products average molecular weight of unconverted fresh feed, gm/gm mol average molecular weight of jet fuel product, gm/gm mol number of experimental observations N R gas constant, kcal/g-mol/OR s, liquid hourly space velocity (cc total feed/hr)/(cc of reactor) T average catalyst temperature, OR residence time, hr t X apparent conversion of feed to products, cc fresh feed/ cc total feed XO jet fuel in fresh feed, LV % Jo
GRI:I.:K
LETTERS
a = preexponential factor for reaction with rate constant
k , hr-I a, = preexponential factor for reaction with rate constant
k , . hr-I A = determinant criterion for multiresponse parameter
estimation v i = stoichiometric coefficient for reaction i pF = PJ =
density of fresh feed, g/cc density of jet product, g/cc
Literature Cited
Froment, G. F., Bischoff, K. B., Chem. Eng. S a . , 16, 189 (1961). Froment, G. F., Bischoff, K. B., ibid., 17, 10.5 (1962). Kittrell, J. It., Advan. Chem. Eng., 8, 98 (1970). Szepe, S.,Levenspiel, O., European Symposium on Chem. Engr. ITr, Brussels, Belgium, September 9-11, 1968. Weekman, V. W., Ind. Eng. Chem. Process Des. Develop., 8, 385 (1969). RECEIVEU for review September 11, 1970 ACCEPTED July 23, 1971
Analytical Solutions for Predicting Vapor-Gas Behavior in Cooler-Condensers James 1. Schrodt Chemical Engineering Department, Cnivarsity of Kentucky, Lexington, Ky. 40506
I n a cooler-condenser, vapor and sensible heat are transferred from the vapor-rioncoiideiisahle gas mixture to the cooled surface a t rates equal to the respective products of the transfer coefficients and driving forces. In the classic design procedure developed by Colburn and Hougen (1934), it is assumed the mixture remains saturated throughout the process, implying that the transfer rates are relatively equal. This situation may be true for some mixtures-e.g., water vapor in air-but most others will teiid either to superheat or subcool and form fog. 20
Ind. Eng. Chem. Process Des. Develop., Vol. 1 1 , No. 1 , 1972
Some researchers have theorized that t'hese tendencies stem from the basic transport characteristics of the mixtures. Colburn and Edison (1941) discussed the various systems which can be expected to superheat or subcool, emphasizing that heavy, sloaly diffusing vapors, such as benzene, butyl alcohol, and toluene, in air are likely to form fog in coolercondensers, whereas light,er, faster diffusing systems like water-carbon dioxide and water-helium should tend to superheat. The C"M was extended by Colburn and Drew (1937)
