Ind. Eng. Chem. Process Des. Dev. 1985, 2 4 , 590-607
598
R,, = reflux ratio at maximum vapor rate (see eq 1) Rmin= minimum reflux ratio ROE= optimum operating reflux ratio S - separation factor = [ x D ( I - x B ) ] / [ x B ( l - X D ) ] V* = design vapor rate = (R* + 1)D* V,, = maximum vapor rate n = reactor conversion X B = bottom composition XD = distillate composition XF = feed composition X F , = ~ minimum ~ ~ operable feed composition XSOR = start of run conversion
erating cost curve is quite linear. Thus, the time-average operating costs for this column should be nearly equal to the operating cost calculated at the time-average feed compition. For this case, little overdesign on the number of trays would be warranted. From Figure 4, the increase in reflux ratio needed to accommodate the full range of feed compositions (xFm = 0.543 and XF* = 0.622) is about 10%. The required overdesign factor for vapor capacity can then be calculated from eq 1.
Greek Letters
or
cy
Vm,/V*
= relative volatility
0~ = reactor residence time 80 = constant = reactor volume/fresh feed flow rate
= 1.07
Nomenclature catalyst activity B = bottom flowrate C, = concentration of reactant A D = distillate flow rate k = reaction rate constant kD = catalyst deactivation rate constant NR = number of theoretical trays in the rectifying section Ns = number of theoretical trays in the stripping reaction NT = number of theoretical trays n = capital cost factor exponent q = feed quality R* = design reflux ratio
Literature Cited
a =
Klttrell, J. R.; Watson, C. C. Chem. Eng. Prog. 1966, 62, 79. Mallk, R. K.; Huges R. R. Comput. Chem. Eng. 1979, 3 , 473. Nishida, N.; Ichikawa, A.; Tazaki, E. Ind. Eng. Chem. Process Des. D e v . 1974, 13, 209. Nlshida, N.; Ichikawa, A,; Tazaki, E. AIChEJ. 1974, 18, 561. Peters, M. S.; Tlmmerhaus, K. D. "Process Design and Economics for Chemical Englneers"; McOraw-Hill: New York, 1980. Smoker, E. H. Trans. AIChE 1938, 3 4 , 165. Swaney, R. E.; Grossmann, I . E. "A Metrlc for Operational Flexiblllty in Chemical Process D d g n " , paper presented at the AIChE Annual Meeting, Los Angeles, Nov 1982.
Received f o r review August 26, 1983 Accepted August 13, 1984
Effect of Phase Behavior on Hydrotreater Performance: Simulation and Experimental VerOf ication Ramakrlshna V. Nailiham, James A. Guln;
Arthur R. Tarrer, and Chrlstlne W. Curtls
Auburn Coal Conversion Laboratory, DepaHment of Chemical Engineering, Auburn Unhersity, Auburn University, Alabama 36849
The effect of vaporization of the liquid feed components on reactor performance for hydrogenation of naphthalene, dissolved in a diluting solvent, is studied. The vapor-liquid equilibrium constant of the vehicle solvent, hydrogen flow rate, temperature, pressure, and feed concentration are shown to affect the naphthalene conversion in several diverse ways. The role of vapor-liquid equilibrium in evaluating hydrotreating kinetics is examined through a simulation model. Experimental results are obtained with a continuous stirred tank reactor for the catalytic hydrogenation of naphthalene in the presence of two vehicle solvents, having considerably different volatilities. The experimental data are in general agreement with the simulation predictions.
Introduction Interest in the catalytic hydrotreatment of Solvent Refined Cod (SRC) and coal-derived liquids has recently increased for two main reasons. First, SRC is a solid fuel and must, therefore, be further processed for use as a transportation fuel. This can be achieved by hydrotreating SRC selectively for maximum production of middle distillates. Second, the deteriorated recycle solvent must be upgraded in order to maintain solvent quality and sufficiency in the process. Thermal hydrogenation of the deteriorated solvent results in only a marginal improvement in solvent quality (Moniz et al., 1983). Catalytic hydrotreatment is, then, an attractive route for selectively upgrading of the recycle solvent. Solvent hydrotreatment can be performed either in a separate reactor or in the same reactor in which SRC is upgraded. The Exxon Donor Solvent process employed the former route while the In0796-4305/05/7 724-0590$07.50/0
tegrated Two-State Liquefaction (ITSL) process at Wils o n d e , AL,employs the latter route. Interest in catalytic hydrotreatment of heavy petroleum resid fractions and heavy crudes has also intensified in recent years due to the decreased availability of lighter crudes and more stringent environmental regulations. Hydrotreating reactions generally occur at higher temperatures (700 to 850 O F ) . In addition, high hydrogen flow rates are used to provide good mixing in the reactor as well as to suppress coke formation on the catalyst through an increase in hydrogen paxtial pressure in the reactor. These operating conditiqns provide a significant opportunity for the components'in the reactor to distribute between the vapor and liquid phases. The extent of vaporization can be as high as 40% under the normal operating conditions in a typical coal liquefaction process (Brunson, 1979). As shown in this paper, the relative distribution of the feed 0 1985 American Chemical Society
Ind. Eng. Chem. Process Des. D e v . , Vol. 24, No. 3, 1985 599 BPR FCV
= = =
BACK PRESSURE REGULATOR FLOW CONTROL VALVE FORWARD PRESSURE REGULATOR PI PRESSURE INDICATOR TI TEMPERATURE INDICATOR CV CHECK VA VE SRVZ SAFETY R ~ VALVE F NC = NORMALLY CLOSED
FPR
$SRV
r”lFEED
u
SZARAAORS
TANK
A X SRV lih
cv
8i
FEED PUMP
I
I:
U
HYOROGEN CYLINDER
U
HYDROGEN CYLINDER
U
HYDROGEN CYLINDER
Figure 1. Continuous hydrotreating unit.
components between the two phases can significantly affect the reactor performance. Singh and Carr (1983) simulated the effect of evaporation of solvent on oil production and total coal conversion in the SRC-I1reactor. Their results indicate that higher oil yields and coal conversion are obtained when a light solvent is used. They attribute this result to the accumulation of solids in the reador leading to higher reaction rates. Brunson (1979) reported that the mean residence time of coal-oil slurry in the reactor can differ from the nominal residence time by as much as a factor of 2 due to solvent vaporization. A similar effect was reported by Gopal et al. (1983). Hydrotreater feedstock is generally not a single compound, but rather contains a large number of components. Of course, all the components are not equally reactive. The purpose of this work is to examine both analytically and experimentally the effects of feed vaporization on hydrotreater performance. Consideration of vapor-liquid equilibrium is important in bench-scale catalyst evaluation and testing and in the design, simulation, and optimization of the hydrotreater. For the work performed here, a naphthalene hydrogenation reaction was selected as a model reaction system. This reaction has considerable relevance to the actual processing of coal liquids since tetralin and naphthalene are contained in the process solvent. A mathematical model is developed for catalytic hydrogenation of naphthalene in a continuous stirred tank reactor (CSTR). Unlike most of the previous studies, liquid vaporization is taken into account in this model by introducing the vapor-liquid equilibrium relations for the respective components. When this model is used, the reactor is then simulated to examine the effects of important process variables such as feed volatility, hydrogen flow rate, temperature, pressure, and feed concentration. The role of vapor-liquid equilibrium in evaluation of apparent hydrotreating kinetics is also examined in this study. Finally, experiments are performed in a continuous hydrotreating unit to verify the key results of simulation studies. Experimental Equipment Batch Reaction System. Batch experiments were conducted in a tubing-bomb microreactor constructed of 316 stainless steel tubing, 1.27 cm 0.d. with 0.071 cm wall
thickness and a volume of 15 cm3. The small tubing-bomb reactor is charged and pressurized with H2, immersed in a temperature controlled fluidized sand bath, and agitated vertically at 860 rpm. Further details of the reaction equipment are given elsewhere (Curtis et al., 1983). Continuous Hydrotreater System. A schematic diagram of the continuous reactor system is shown in Figure 1. The reactor is a spinning-basket continuous stirred tank (CSTR) reactor. Ebullated-bed reactors are generally used for hydrotreatment of coal derived materials; however, the mixing pattern in ebullated-bed reactors closely approximates that in a CSTR (Bickel and Thomas, 1982). The main advantage of this reactor is that gradientless operation can be achieved easily (Tajbl et al., 1966). A 300-cm3autoclave designed and built by Autoclave Engineers serves as the basis for the reactor system and was equipped with a packless magnetic drive to rotate the catalyst basket. To reduce the liquid holdup in the reador, a 2.54 cm i.d., 4.44 cm o.d., and 11.43 cm long stainless steel sleeve was inserted into the autoclave. The catalyst was held in a perforated annular basket, made of 0.046 cm thick 304 stainless steel plate with 0.084 cm diameter perforations. An impeller of 2.22 cm diameter was attached to the bottom of the agitator shaft below the basket to provide good mixing. The liquid level was maintained by means of a 0.318 cm diameter dip-tube. Mixing and liquid level control were evaluated in a plexiglass model of the reactor at ambient temperature and 20 psig pressure. A high degree of gas entrainment was observed in the liquid phase and the liquid level was maintained constant below the tip of the dip-tube. A thermocouple was placed above the top of the catalyst basket to measure the liquid temperature. The reactor was heated by means of an electric furnace and the temperature was controlled by means of a Barber-Coleman dual instrument controller. Hydrogen gas was metered to the reactor from the high-pressure cylinder by means of a Hastings mass flow meter. Liquid feed was pumped to the reactor by a Ruska high-pressure metering pump. The pump accurately delivered the set flow rate independent of the downstream pressure. The combined gas and liquid feedstream entered the reactor at the bottom. The liquid and gas product from the reactor was cooled in a double pipe heat exchanger to ambient temperature before it entered the gas-liquid separation system. Liquid
600
Ind. Eng. Chem. Process Des. Dev.. Vol. 24, No. 3, 1985
Table I. Catalyst Properties and Presulfidina Conditions catalyst Harshaw 0402T size 0.318 X 0.318 cm extrudates 3% COO and 15% Moo3 composition surface area 200 m2/g pore volume 0.4 cm3/g reactor type H,S concentration in H, reactor pressure catalyst charge temperature program 200 "C 315 "C 370 "C
L
packed bed 5%
atmospheric 10 !. for 2 h for 1 h for 2 h
Table 11. Reaction Conditions for Batch and Continuous Experiments Batch Experiments catalyst charge 0.125 g liquid charge 5g concentration of naphthalene 2 wt % in the liquid charge agitation rate 860 rpm reactor pressure 50, 82, 150, 200 atm reaction temperature 250, 300, 350 "C Continuous Experiments catalyst charge to the basket 1.6 g reactor pressure 82 atm reaction temperature 300 oc stirrer speed 1000 rpm 10 wt % concentration of naphthalene in the liquid feed H2gas feed rate 100, 250, 500 sccm 30 cm3 liquid holdup
sampling ports were provided immediately after the cooler and the first-stage separator. The liquid product from the separators was collected in a 1-L autoclave and emptied periodically without disturbing the system pressure. The gases were passed through a Grove back-pressure regulator and a wet-test meter before venting to the atmosphere. Materials. The following chemicals were used as received: cyclohexane (Fisher, certified ACS), white oil (Thomas Chemicals), naphthalene (Fisher, scintanalysed), and hexadecane (Aldrich, 99%). White oil met the standard USP20 specifications. The properties of the white oil were estimated by the methods given in Maxwell (1955). Hydrogen gas cylinders (2500 psi, commercial grade) were supplied by Alabama Oxygen Co. The catalyst used was Harshaw 0402T cobalt molybdenum supported on A1203,0.318-cm extrudates. The properties of the catalyst are listed in Table I. The catalyst was presulfided in a packed bed using a mixture of 5% H2S and 95% H2 at the conditions listed in Table I.
Experimental Procedure Batch Reactor. The reaction conditions for the batch experiments are given in Table 11. In a batch run, the tubing bomb was charged with the liquid reactant and presulfided catalyst pellets. The reactor was pressure tested and then charged with hydrogen at ambient temperature. After being charged with hydrogen, the tubing bomb was attached to the vertical shaft on the agitation assembly and agitated at ambient temperature for 5 min at 860 rpm and then submerged into the preheated fluidized sand bath. The heat-up time was about 1min. At the end of the specific reaction time, the tubing bomb was removed from the sand bath, immediately quenched in tap water, and checked for any leaks. The gases were released from
Reactor
Condenser-Cooler
Figure 2. Schematic diagram of the hydrotreater system.
the bomb and the liquid product was analyzed by gas chromatography with a Varian 3700 equipped with a SP2250 column (2.4 m X 0.3 cm o.d.), FID detection, and temperature programming from 60 to 200 "C/min. Continuous Reactor. The reaction conditions for the continuous experiments are given also in Table 11. In a continuous run, the catalyst basket was uniformly filled with catalyst pellets and glass beads and was sealed at the top by the insertion of a concentric circular perforated plate in the free space of the basket. The plate was secured in place by means of a stainless steel wire. The reactor was sealed and the entire unit was pressure tested. When the hydrogen flow was stabilized at the desired flow rate, the liquid feed pump was started. After the reactor was filled with a sufficient amount of liquid, the furnace was turned on, and the reactor was heated to the reaction temperature at the rate of about 100 "C/h. Steady state was assumed to be reached when the product concentrations were constant over an 8-h period. Liquid samples were collected after every hour, and the samples were analyzed by gas chromatography equipped with a SE30 column (2.4 m X 0.3 cm o.d.), FID detection, and temperature programming from 80 to 270 "C at the rate of 8 "C/min. All the continuous runs were performed on the same catalyst charge. The last flow run was a repeat of the first flow run, using a naphthalene-white oil feed system to check for catalyst deactivation. No significant catalyst deactivation was detected. About 50 ppm of CS2 was added to the feed liquid to maintain a slight H,S partial pressure in the reactor for maintenance of catalyst activity and stability. Hydrotreater Model. Basically, the model consists of a set of material balance equations for the respective components. Since the reactor operates nearly isothermally, heat balance constraints are not considered in the model development. The model is based on the following assumptions: (1)There is complete mixing in the reactor. (2) The gas and liquid are in equilibrium within the reactor. (3) The vapor-liquid equilibrium constants are functions of only temperature and pressure. (4) The homogeneous reaction rate is negligible. Assumption (4)has been verified in batch reaction experiments in this laboratory (Nalitham, 1983). The other assumptions are discussed later in this paper. Figure 2 shows a schematic diagram with appropriate variables for the hydrotreater model. The vapor and liquid streams from the reactor pass through a condenser and into a separator. Based on the previous assumptions, the following material balance equations can be written over the reactor
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
601
In the above equations the subscript the i may refer to naphthalene (N), tetralin (T),solvent (S),or hydrogen (H). Note that the solvent does not participate in the reaction and, therefore, as in eq 1 is zero.
Results and Discussion Naphthalene Hydrogenation Kinetics. A kinetic rate expression for naphthalene hydrogenation was developed from batch experiments for use in the continuous hydrotreater studies. Under the reaction conditions studied, tetralin was the only reaction product from naphthalene. The reaction scheme is thus represented as f
2H2
*
naphthalene
a
0.I
0.0
0.4
tetralin
The completely hydrogenated product decalin was not detected under the conditions of this study. Second-order kinetics were used to fit the batch data. Strictly speaking, the rate expression should include adsorption terms for naphthalene, tetralin, hydrogen, and solvent. However, the experimental data were observed to fit the simple power-law model reasonably well. Validity of the second-order kinetics for naphthalene hydrogenation has also been verified by other workers (Sapre and Gates, 1981; Shridharani, 1983). The rate expression can be written as where k is the rate constant and CN and CHI are the concentrations of naphthalene and hydrogen in the liquid phase, respectively. In each batch experiment, an excess amount of hydrogen was charged to the reactor. The ratio of the amount of hydrogen available in the reactor to the amount of hydrogen required to convert all the naphthalene completely to tetralin was about 20. Thus, for all practical purposes, the partial pressure changes due to consumption of hydrogen in the reaction may be neglected. The solubility of hydrogen in the liquid medium is mainly a function of temperature and pressure (Chao et al., 1980). At constant temperature and hydrogen partial pressure, eq 5 can be written as a pseudo-first-order reaction where
k ’ = kCHl
0.0
I.2
I.6
2.0
R E A C T I O N TIME, hr.
(7)
The rate parameter k’ was determined from the slope of a CNIvs. t plot on a semilogarithmic scale. Figure 3 shows a typical plot of the experimental data at constant temperature and pressure. The rate parameter, k’, was determined at four levels of reactor pressure: 50,82,150 and 200 atm. Then a linear regression of k’ with CHI was used to estimate the rate parameter, k . This procedure was repeated for three different temperatures; 250, 300, and 350 O C , to estimate the activation energy and frequency factor. The rate constant so determined is listed in Table 111. The effect of agitation rate on conversion was studied at three levels of agitation; 600,860, and lo00 rpm, to check
Figure 3. Typical plot of batch kinetic data. Table 111. Reaction Conditions and Parameters Used in Simulation Baseline Reaction Conditions temperature 300 DC pressure 100 atm
k = 2.12
X
Rate Constant lo5 exp(-9171.3/T) L*/(gmol g h)
Temperature Dependency of K Values a t a Pressure of 100 atm Ki = exp(Ai + Bi/T) hydrogen hexadecane hydrogen toluene tetralin
Ai -0.4489 9.0201 -3.0085 6.5468 6.6305
Bi 1121.92 -7345.64 2610.0 4082.69 -5168.74
Pressure DeDendencv of K Values a t a TemDerature of 300 OC 50 atm 100atm 150 atm 200atm 250atm hydrogen 8.75 4.86 3.59 2.93 2.55 hexadecane 0.0465 0.0286 0.0242 0.023 0.0216 hydrogen 7.26 4.03 2.76 2.07 1.56 toluene 0.785 0.588 0.559 0.574 0.636 tetralin 0.1736 0.1017 0.0767 0.0667 0.0608
hydrogen naphthalene tetralin
Liquid Molar Volumes L/g-mol 0.03 toluene 0.151 hexadecane 0.177
L/g-mol 0.170 0.372
for mass transfer limitations. The conversion was the same at all agitation rates. Reactor Simulation. The continuous CSTR was simulated for two hypothetical feed compositions to study the effect of the phase behavior on naphthalene conversion. Feed 1 contained naphthalene and an aliphatic solvent component, hexadecane. Feed 2 contained naphthalene and an aromatic solvent, toluene. The reaction conditions and the parameters used for the simulation are listed in Table 111. Wherever feasible, experimental values were used as parameters in the simulation. In this respect, the simulation representa an actual experiment. As discussed
602
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
earlier, the rate constants were determined experimentally from batch data and expressed as a function of temperature in the form of a classical Arrhenius relation (Table 111). Vapor-liquid equilibrium data, reported by Chao et al. (1980),on the binary systems tetralin-hydrogen, toluene-hydrogen, and hexadecane-hydrogen were used in this work. The K values were shown to be strong functions of temperature and pressure and weak function of composition. As a first approximation,K values are assumed to be independent of composition of this work. Experimental K values for naphthalene at the simulation conditions are not available in the literature. As an approximation,they are estimated by assuming that the K values for naphthalene and tetralin are in proportion to their vapor pressures. Using the mathematical model developed in the previous section, the effects of temperature, pressure, hydrogen flow rate, and feed composition on conversion can be simulated for the two model feedstocks. Material balance equations 1 to 4 together with the rate expression (eq 5) describe the reactor behavior. The molar concentration, Cn, appearing in eq 5 is related to the mole fraction x i by the following equation
11.
Toluene, 100% Encesr H p
111.
Hoxadecana. 4 0 0 % Excerr IK, * 0 . 0 2 8 6 )
IV.
Hexadocane, 100% Excerr H z
Hp
24
40
60
80
IO
SPACE VELOCITY, cc/hr-g
The partial molar volume, Vi, appearing in the above equation is a function of temperature, pressure, and composition. These data are not generally available for multicomponent systems and therefore as an approximation, pure liquid molar volumes are used for the partial molar volumes. Since the critical temperature of hydrogen is considerably lower than the reaction temperature, the pure liquid molar volume for hydrogen is estimated at its critical temperature and the reactor pressure. The liquid density can be computed from eq 8 using the pure liquid molar volumes according to
CMixiR v~xiR
Toluene, 4 0 0 % Excesr H Z IK, = 0.5881
Feed1 I O % Naphthalrne
20
Xi cfi= c vixi
P=-
I.
(9)
Utilization of these approximations, based on ideal solution behavior, results in computed liquid density values at representative liquid compositions which lie between the liquid solvent density and the liquid naphthalene density. In performing the simulation, eq 1to 4 along with eq 5 and 8 are solved for the unknown quantities Fm FvR,xiR, and ylR using the known quantities Fo, zi, Ki, k , and Vi. The solution was obtained by the secant method for simultaneous nonlinear equations (Phillip, 1959). The naphthalene conversion is calculated according to
Naphthalene conversions are calculated at various levels of liquid feed space velocity defined as
where the sum is taken only over N, T, and S. Effects of Hydrogen Flow Rate on Conversion. Equations 1 to 4 along with eq 5 and 6 have been solved for various values of hydrogen flow rates for the two liquid feeds. The results are plotted in Figure 4 as naphthalene conversion vs. liquid space velocity. In Figure 4, the hydrogen flow rate is shown as the percentage of excess hydrogen, which is defined as the normalized amount of hydrogen supplied in excess of the theoretical amount required to convert naphthalene completely to decalin.
Figure 4. Effects of solvent volatility and hydrogen flow rate on hydrotreater performance.
Mathematically, the percentage excess Hzis defined by the following equation.
A comparison of the results for the two feed compositions reveals several interesting points. As expected, naphthalene conversion decreased with increasing liquid space velocity for both feeds. For a given space velocity and for a given amount of excess hydrogen, the naphthalene conversion with the naphthalene-toluene (NT) feed system is higher than that with the naphthalene-hexadecane (NH) feed system. The differences in conversion can be rationalized as follows. The K value for toluene (0.588) is about 20 times greater than the K value for hexadecane (0.0286). The large K value for toluene indicates that most of the toluene is in the vapor phase in the reactor and that the reactor liquid contains mainly naphthalene, tetralin, and hydrogen. For example, at a space velocity of 20 cm3/(h g) and with 400% excess hydrogen, the concentrations of naphthalene and hydrogen in the liquid phase are 0.44 and 1.07 g-mol/L, respectively, with the NT system compared to the corresponding values of 0.27 and 0.71g-mol/L with the NH system. According to the kinetic rate equation (5), the rate is proportional to the concentration of hydrogen and naphthalene in the liquid phase, and thus higher reaction rates are obtained with the N T system. As the amount of vaporization increases, e.g., with toluene as the solvent, the mean residence time of the remaining liquid is increased. Combination of these factors leads to higher conversions of naphthalene using the more volatile solvent. It should be noted that for the NH system, the hydrogen flow rate has little effect on conversion. In fact, the conversion drops slightly when the hydrogen flow rate is increased. Hexadecane is relatively nonvolatile, and, as the hydrogen flow rate is increased, the amount of naphthalene vaporized is greater than the amount of hexadecane vaporized. Naphthalene in the vapor phase is not accessible for reaction on the catalyst surface. In effect, this results in a decrease in naphthalene conversion with an increasing
Ind. Eng. Chem. 96
1
I\ I. 11.
Process Des. Dev., Vol. 24, No. 3, 1985 603
I.
Toluene, 4 0 0 % Excess Hp
11.
Toluene. 100% Excess H 2
IV.
Hexadecane. 100% Excess Hg
Toluene. IO% Naphthalene Toluene. 2 % Naphlhalena
111.
Hexadecane. 10% Naphthalene
IV.
Hexadecana, 2% Navhthalene
8 z* 5 5
9
Space V ~ l o c l l y20cc/hr-g ~ Feeds I O % Naphthalene
cn [L
W
> z
45
0 0 W
35
2 I
I
\\\
-J
Q
I
I-
I LL
25
Q
z 15
5 220
SPACE VELOCITY, c c / h r - g
240
260
280
300
320
TEMPERATURE, ' C
Figure 5. Effects of solvent volatility and feed concentration on hydrotreater performance.
Figure 6. Effects of solvent volatility and temperature on hydrotreater performance.
amount of hydrogen t~ the reactor. In the case of the NT feed system, the converse is true; i.e., conversion increases with increasing hydrogen flow rate. This can be understood by noting that the hydrogen flow rate has two opposing effects on the reaction rate one effect is an increase in conversion due to solvent vaporization and the second effect is a decrease in conversion due to reactant (naphthalene) vaporization. For a light solvent, the former effect dominates the latter effect. The net result is an increase in conversion with an increasing hydrogen flow rate for the NT system. Effect of Feed Composition on Conversion. The effect on reactor performance of naphthalene concentration in the liquid feed was studied with the two feed systems for a constant value of percent excess hydrogen (eq 12). The results are shown in Figure 5. With the NH system there is little effect on conversion when the liquid feed concentration is increased from 2 w t % to 10 w t % naphthalene. When the naphthalene concentration in the feed was increased, the hydrogen flow rate was increased correspondingly to keep the ratio of inlet feed concentrations a t a constant value. A higher hydrogen flow rate resulted in an increase in vaporization of toluene and, in light of the reasons presented earlier, higher conversions were obtained. With 2 w t % naphthalene in the liquid feed, the differences in conversion were small (curves I1 and IV) for the two feed systems because the hydrogen flow rate is low. Generally speaking, as the hydrogen feed rate becomes smaller, the effecta caused by differences in solvent volatility become smaller as well. Effect of Temperature. Temperature affects the kinetic rate constant as well as the vapor-liquid equilibrium constants (K values). The temperature dependency of the rate constant is described by the classical Arrhenius relation as shown in Table 111. The K values for various components at constant pressure were fitted by an equation of the form Ki = exp(Ai B i / T ) (13) Experimental data reported by Chao et al. (1980) were used to estimate the constants Ai and Bi in eq 13 for tetralin, toluene, hexadecane, and hydrogen. Estimated
values of Ai and Bi are listed in Table I11 for a pressure of 100 atm. ?she K value for naphthalene was computed from'the K value of tetralin according to their vapor pressure ratio. Figure 6 shows the effect of temperature on reactor performance for the two feed systems. At low temperatures, the differences in conversion for the two feed systems are small because the amount of liquid feed vaporized is less than that at higher temperatures. As the temperature is increased, the amount of liquid vaporization increases and the differences in conversion become greater. It should be noted again that with the NH system, the hydrogen flow rate has little effect on conversion whereas with the NT system, the conversion increases with hydrogen flow rate. The effect of neglecting liquid-phase vaporization upon determination of reaction kinetics rate constants was also studied by use of the simulation program. To do this, an apparent reaction rate constant at various temperatures was computed using the previously generated conversion data in Figure 6 together with eq 1-5 assuming yIR = 0 for i = N, T, and S, Le., neglecting all evaporation of the liquid phase components. These apparent rate constants are plotted in an Arrhenius plot as shown in Figure 7. For the NH system, the computed data lie on a straight line. However, for the NT system, the plot is slightly curved, an effect caused by the neglect of vaporization. Activation energies were computed with the data in Figure 7 by linear regression. The computed apparent activation energies were different from the actual activation energies (Table 111) used in generating the data in Figure 6. The apparent activation energy for the NH system was 20 kcal/mol, which was not very different from the actual value of 18.4 kcal/mol But the apparent activation energy for the N T system was 29.8 kcal/mol, which was about 50% greater than the actual activation energy. Also the apparent frequency factor for the NT system was greater than the actual value by a factor of 6 X lo4. Thus, the failure to properly account for phase behavior in kinetic analysis leads to falsification of activation energies and frequency factors. Similarly, it may be expected that the concentration dependence (order of reaction) would be distorted if liquid vaporization effects were not taken into
+
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
604
0. I
1 'molal
t-
z
a
I-
0.01
v,
-
z
0 0 W
-I
a [L
tz w [L
6 0.001
a a Q
0.000 I I 7
I .8
I .9
I /T,
(
2.0
I /K) x I
2.2
2 I
o3
0
v)
[L
w
60
0 0
w z w
50
i = l
I c
I G 40
a z
30
2 01
50
IO0
150
200
20
30
40
50
Figure 9. Effect of solvent volatility on hydrotreater performance: experimental results.
9
> z
10
SPACE V E L O C I T Y , c c / h r - g
Figure 7. Arrhenius plot of apparent rate constants.
1
o
250
1
300
R E A C T O R PRESSURE, atm.
Figure 8, Effects of solvent volatility and pressure on hydrotreater performance.
account in the kinetic analysis. This analysis has strong analogies to the classical diffusion-influenced kinetic analysis as embodied in the effectiveness factor concept of catalysis. Effect of Reactor Pressure. In general, K values decrease with increasing pressure. Experimental K values (Chao et al., 1980) at different pressures are given in Table I11 for a temperature of 300 "C. The K value of naphthalene again was computed from that of tetralin according to the vapor pressure ratio. The effect of reactor pressure on conversion is shown in Figure 8; the conversion increased with increasing pressure for both feed systems. However, the sensitivity of conversion to reactor pressure for the NT system is lower than that with the NH system at an excess hydrogen level of 400%. With the NH system, an increase in pressure causes the concentration of naph-
thalene and hydrogen in the liquid phase to increase and the mean residence time of the liquid to decrease. However, the former factor dominates the latter factor resulting in an increase in conversion with increasing pressure. Recall that the rate constant is proportional to hydrogen concentration. With the NT system, however, an increase in pressure causes more toluene to be in the liquid phase, and thus the reactor liquid is diluted with solvent. The solvent dilution effect partly offsets the increase in reactant concentrations and conversion. It should be noted that at low hydrogen flow rates (100% excess hydrogen), the amount of liquid vaporization is small, and, hence, the sensitivity is nearly the same for both feed systems. Experimental Verification of Simulation Results The key result of the simulation studies was that higher naphthalene conversions were obtained when a more volatile vehicle solvent was used in preparing the hydrotreater feed. To verify this result, continuous reactor experiments were performed with two feedstocks. The two feedstocks were 10 wt % naphthalene in cyclohexane (NC) and in white oil (NW) and 400% excess H2 was used. Cyclohexane is a more volatile solvent compared to white oil. The experimental results are shown in Figure 9. In order to determine the naphthalene conversion, liquid samples from the separator (Figure 2) were analyzed. Since the separator was operated at ambient temperature, the gas leaving the separator was considered to be pure hydrogen (YH= 1). Naphthalene conversion was computed from the following material balance equation, equivalent to eq 10
A comparison of the curves for the two feedstocks shows that for a given space velocity, higher naphthalene conversions are obtained with the NC system than with the NW system. The question then arises as to why the conversions differ with the two solvents. Several possibilities exist to explain these differences. As shown by the simulation results, one possibility is the differences between
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
005
I .2
W
z w
SPACE VELOCITY- 23.I eelhr'l DILEYT SOLVENT- CYCLOHLXAU
50
-I
Q
I
n
t I
,..
00 2
3
4
A,
5
DIMENSIONLESS TIME,
6
LL
h l
a
7
e
Figure 10. Residence time distribution of liquid in the reactor.
the K values of the two solventa, provided the assumptions made in the model development are valid in the present analysis. The validity of these assumptions is examined in the following paragraphs together with various alternative explanations for the observed experimental results. Mixing Pattern in the Reactor. In order to use eq 1, the spinning basket reactor should approximate a single perfectly mixed reactor. To check this assumption, the residence time distribution of the liquid was evaluated using an inert-pulse technique (Levenspiel, 1972). A pulse of tetralin was injected into a pure.hexadecane feed. The experimental dimensionless exit age distribution EBat different levels of agitation and the predicted distribution for a perfectly mixed reactor are plotted in Figure 10. Except for the small increase in the exit concentration during the initial period, which is probably due to slight bypassing, the exit age distribution follows an exponential curve at all the three levels of agitation and approximates the exit age distribution expected for an ideal CSTR. Interphase Mass Transfer Resistances. According to the film theory, the following mass transfer processes are operative: (i) mass transfer through the gas film (ii) mass transfer through the liquid film at the gas-liquid interface
r1 = kP&Cn. - Cl)
40
Z
(16)
(iii) mass transfer through the liquid film at the liquidsolid interface
r* = k,a,(G - C,) (17) At steady state, mass transfer rates are equal to the reaction rate (eq 5). The concentration gradients in eq 15, 16, and 17 can be estimated from the experimental rate values and the mass transfer coefficients. Since the gas phase contains mostly hydrogen gas and the mass transfer coefficient is large relative to the liquid side coefficient, the concentration gradient across the gas film (eq 15) is negligible. The liquid side gas-liquid mass transfer coefficient and solid-liquid mass transfer coefficient are esimated using the correlations of Akita-Yoshida (1973) and Sano et al. (1974), respectively. The estimated concentration differences (eq 16 and 17) were less than 5% of the actual concentrations for the corresppnding components in the liquid phase. The effect of stirrer speed on conversion is shown in Figure 11. The conversion profile is flat, indicating that interphase mass transfer
30
20 30
60
90
120
I50
180
AGITATION RATE, rpm x I o-'
Figure 11. Effect of agitation rate on hydrotreater performance.
resistances are probably negligible in the present work. Equilibium between Gas and Liquid. Wilson et al. (1981) reported vapor-liquid equilibrium studies using aromatic compounds in which equilibrium was reached even when the residence time in the flow cell was less than 10 s. In the present study, the residence time of both gas and liquid is of the order of several minutes, and thus assumption (2) is reasonably valid. Catalyst Poisoning. It is well-known that a hydrotreating catalyst can suffer a decline in intrinsic activity due to adsorption of impurities such as basic nitrogen compounds and phenols on its surface (Gates et al., 1978; Nalitham, 1983). The lower values of naphthalene conversion observed with the NW feed system could be the result of inhibition by impurities present in the white oil. For example, Satterfield and Way (1972) reported studies of the isomerization of cyclopropane on a silica-alumina catlayst in a trickle-bed reactor using different nonvolatile inert liquids in which the rate of reaction was inhibited by the aromatic species in the paraffinic liquid. To check for this effect, an experiment was conducted in which the feed to the reactor was switched from NW feed to NC feed. Steady state was essentially reached within a period equal to about six time constant values (about 4 h) as determined by consistent analytical results. If the catalyst had retained strongly adsorbed poisons, steady-state attainment with respect to the catalyst would not have been possible within that short period. To verify this concept, another experiment was conducted in which the catalyst was poisoned with quinoline, a basic nitrogen compound. After the steady state, at a lower level of activity, was reached, the feed to the reactor was switched from poison (quinoline)-containing feed to a poison-free feed. In this experiment, steady state, at a higher level of activity, was reached only after about 2 days (60 time constant values). To provide a further check on the absence of inhibition effects with the NW feed system, one experimental run was made with a naphthalene-hexadecane feed system (NH). The hexadecane lot used in this run was highly refined (Aldrich, 99% pure) and, hopefully, free of impurities. The volatility of hexadecane is comparable to that of white oil. Naphthalene conversions were nearly iden-
606
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
Table IV. ExDerimental Results and Fitted K Values for NaDhthalene-CYclohexne and Nauhthalene-White Oil Systems results at each space velocity, cm3/(h g) white oil cyclohexane 9.2 23.1 46.1 9.2 23.1 46.1 naphthalene conversion 34 25 18.5 67 46 30 fitted K values solvent hydrogen naphthalene tetralin
0.011 10.60 0.095 0.115
0.010 7.35 0.086 0.112
0.011 5.75 0.079 0.098
0.479 14.52 0.080 0.090
0.452 12.02 0.091 0.102
0.420 11.40 0.071 0.089
liquid phase concentrations, g-mol/L solvent hydrogen naphthalene tetralin
2.48 0.30 0.28 0.13
2.47 0.45 0.33 0.10
2.45 0.60 0.37 0.08
0.27 4.2 0.59 1.07
0.34 4.4 0.82 0.63
0.37 4.6 0.97 0.36
sum of square of error
0.45 X
1.0 X lo-"
0.36 X
0.31 X lo-"
0.18 X lo-'*
0.19 X
lo-'*
ticaly with both NH and NW systems. If white oil had impurities, which were detrimental to the catalyst activity, the conversion with the NW system would have been lower than the conversion with the NH system. It seems clear from these studies that the lower naphthalene conversions with the NW system were not the result of a catalyst poisoning effect. Solvent Kinetic Effects. The catalytic kinetic rate function can depend not only on temperature but also on the type of vehicle solvent. The differences in rate function values with different solvents are caused by the deviations from ideal solution behavior and the differences in competitive adsorption between the reactant and the inert solvent on the active sites of the catalyst (Eckert, 1967). The average solubility parameter of the solution is an indicator of the departure from ideal solution theory. The solubility parameters are 8.2, 7.99, and 8.03, for cyclohexane, hexadecane, and heptadecane, respectively (Chao and Seader, 1961). Earlier results have shown that naphthalene conversion is not changed when hexadecane is substituted for white oil. White oil is a paraffinic mixture composed mainly of CI8-CMaliphatic compounds. Thus, it seems reabonable that the solubility parameter for white oil is nearly the same as the solubility parameter for hexadecane. Adsorption of cyclohexane on Co-Mo hydrotreating catalyst is negligible (Kim, 1978). White oil is highly paraffinic and prbbably is not adsorbed on the catalyst in significant amounts compared to the aromatics such as benzene, quinoline, and naphthalene (Kipling, 1965). Based on these considerations, the intrinsic kinetic rate constant is assumed to be the same for naphthalene hydrogenation in the presence of cyclohexane and white oil. Effects of Pore Diffusion. Based on experiments by Gollakota (1983) and theoretical calculations following Froment and Bischoff (1979), it is estimated that the effectiveness factor for the catalyst pellets used in this work is greater than 0.8. The production of decalin, a product which would be favored by pore diffusional influences, was not observed in this work. The same catalyst pellets were utilized in both the batch reactions for determination of the reaction rate constants and in the continuous reactor, and any effects of pore diffusion should be essentially accounted for by this procedure.
Effect of Phase Behavior on Conversion Based on the results of the simulation, and the consideration of alternative explanations presented above, the hypothesis advanced in this work is that the differences in naphthglene conversion with the two model feedstocks are due to the differences in the extent of liquid vapori-
zation. The dashed curve in Figure 9 represents the predicted conversion profile, assuming no vaporization of three of the liquid phase components, i.e., YNR = ySR= ym = 0. The dashed curve lies between the two experimental curves. This behavior is consistknt with the model prediction? as shown in the following. Actual experimentally determined K values for the multicomponent systems H2-naphthalene-tetralin-cyclohexane and H2-naphthalene-tetralin-white oil are not known. In order to apply the mathematical model to the experimental results for the NC and NW systems, the equilibrium K values are treated as adjustable parameters for fitting the experimental data. The fitted K values so determined are then compared with K values estimated from the binary data of Chao et al. (1980) to check their validity. An alternative procedure would have been to estimate reasonable K values from data, e.g. (Chao et al., 1980) directly, and then to use these in the simulation, but this procedure is essentially equivalent to the former. Thus, the experimental conversion data are fitted into the mathematical model (eq 1-4) to estimate the unknown K values for the respective components. Since the experimental information is available on X , (see Figure 2)) not on x ~ R , the following additional material balances were written to relate the reactor effluent streams, separator effluent streams, and the experimental conversion. FG, - (FvY,
+ FIX,) = - a i F g Z N X N
CYi
(18)
=1
As mentioned previously, the separator is operated at the ambient temperature, and hence, YN, YT, and ys are set to zero. Equations 1to 5 , 8 , and 18 to 21 are solved for the unknown quantities F, 4,FvR, Fm,yIR,x,, xiR, and Kiusing the known quantities F,, z ~ XN, , k, and V,. The vapor-liquid equilibrium constants obtained at each conversion level and the sum of square of error between the lefbhand side and right-hand side of eq 1to 4 are given in Table IV. An examination of the K values in Table IV indicates that the fitted K values for cyclohexane are about 50 times higher than the corresponding values for white oil. This is consistent with the relative volatility of the two components. Also, the K values for naphthalene and tetralin are about the same in both cyclohexane and white oil and are less than the K value for cyclohexane. The K values for the individual components are slightly different at different conversion levels. This is reasonable
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 3, 1985
Table V. Experimental K Values for Several Binary and Ternary Systems" component K value component K value hydrogen hexadecane hydrogen toluene hydrogen tetralin
6.22 0.0349 5,15 0.657 14.83 0.1269
hydrogen
tetralin diphenylmethane
hydrogen m-xylene
"Conditions: temperature, 300 Chao et al. (1980).
OC;
16.96 0.123 0.054 13.4 0.154 0.383
pressure 82 atm. Source:
in that the K values are also actually weak functions of composition. Chao et al. (1980)measured the K values for several binary systems in a flow apparatus. Experimental K values for representative systems are given in Table V. The agreement between the fitted K values listed in Table IV and the experimental values listed in Table V is considered satisfactory. The concentrations of naphthalene in the liquid phase in the reactor at different space velocities are also listed in Table IV. A comparison of the naphthalene and hydrogen concentrations in the liquid phase for the two solvents show that for a given space velocity, the concentrations are higher with the NC system than with the NW system. The higher naphthalene conversions with the more volatile NC system are due to the factors discussed earlier with regard to the simulation behavior. Conclusions Simulation of the CSTR for hydrogenation of naphthalene, dissolved in a vehicle solvent, has shown that the vapor-liquid equilibrium constants, hydrogen flow rate, feed concentration, temperature, and pressure affect the reactor performance. With a relatively heavy solvent, such as hexadecane, naphthalene conversion is not significantly changed by increasing the hydrogen feed rate or by increasing the liquid feed concentration. However, with a lighter solvent, such as toluene, naphthalene conversion is increased with increasing hydrogen feed rate and increasing liquid feed concentration. With the lighter solvent, the increase in concentration of reactant in the liquid phase and the increase in mean liquid residence time gives higher naphthalene conversions. It is shown that failure to account for liquid vaporization effects in kinetic analysis can lead to false activation energies and frequency factors, a phenomenon having strong analogies to the classical diffusion-limited kinetic analysis. The sensitivity of conversion to reactor pressure is greater for the heavier-feed system. Experimental results with two model feed systems, naphthalene-cyclohexane and naphthalenewhite oil, agree in principle with the simulation results. For a given space velocity, naphthalene conversion with cyclohexane as a solvent is significantly higher than that with white oil as a solvent. Consideration of various possible reasons for this observation, together with the simulation results, shows that phase-behavior phenomena are a sufficient and reasonable explanation of the observed results. Nomenclature a, = liquid-solid interfacial area, cm2/cm3of solid a = gas-liquid interfacial area, cm2/cm3of liquid Cg = concentration, g-mol/L Ci, = concentration of gas-liquid interface, g-mol/L of gas Ca = concentration at gas-liquid interface, g-mol/L of liquid C, = concentration on solid surface, g-mol/L Eo = dimensionless exit concentration
607
F = total molar flow rate,,g-mol/h
K = vapor-liquid equilibrium constant k = rate constant, L2/(h g of cat. g-mol) k,, k,, k1 = mass-transfer coefficients, cm/s k' = pseudo-first-order rate constant, L/(h g of cat.) M = molecular weight P = vapor pressure, atm r = reaction rate, g-mol/(h g of cat.) rg,rl, r, 7 mass transfer rates, g-mol/(s cm3) t = reaction time, h f = average residence time of liquid in the reactor, min = liquid molar volume, L/g-mol V = partial molar volume, L/g-mol W = catalyst charge, g x = liquid mole fraction X = conversion y = vapor mole fraction z = mole fraction in feed Subscripts g = gas
H = hydrogen i = ith component 1 = liquid phase N = naphthalene 0 = inlet R = reactor S = solvent T = tetralin v = vapor phase Greek Letters 6 = dimensionless time ai = stoichiometric coefficient r = space velocity, cm3/(h g of cat.) p = liquid density, g/cm3 Literature Cited A k b , K.; Yoshlda, F. Ind. Eng. Chem. Process Des. Dev. 1073, 12, 78. Bickel, T. C.; Thomas, V. G. Ind. Eng. Chem. Process D e s . D e v . 1082, 21, 377. Brunson, R. J. Fuel 1970, 58, 203. Chao, K. C.; Kim, H. Y.; Ollphant, J. L.; Sebastian, H. M.; Simmick, J. J. "Phase Equilibrium in Coal Liquefaction Processes", Final Report Prepared for Electric Power Research Instltute, Palo Alto, CA, Report No. AP-1593, 1980. Chao, K. C.; Seader, J. D. A I C M J . 1081, 7 , 598. Curtis, C. S.; Guin, J. A.; Tarrer, A. R.; Huang, W. J. Fue/ R o c . Techno/. 1083, 7, 277. Eckert, C. A. Ind. Eng. Chem. 1987, 59, 20. Froment, 0. F.; Blschoff, K. B. "Chemical Reactor Analysis and Deslgn"; Wiley: New York, 1979; p 195. Gates, B. G.; Katzer, J. R.; Olson, J. H.; Kwart, H.; Stiles, A. B. U S . Dept. of Energy Report No. FE-2028-13, 1978. Gopal, J. S.; Shah, Y. T.; Carr, N. L. Can. J . Chem. f n g . 1083, 67, 603. Kipiing, J. J. "Adsorption from Solutlons of Non-Electroiytes"; Academlc Press: New York, 1985. Kim, C. M.S. Thesis, University of Utah, SaR Lake City, UT, 1978. Levenspiel, 0. "Chemical Reaction Engineering"; Wiley: New York, 1972. Maxwell, J. 0. "Data Book on Hydrocarbons". D. Van Nostrand Co., Inc.: New York, 1955. Monk, M.; Davies, 0. L.; Cantrell, C. E. Paper presented at the Elghth Annual EPRI Contractor's Conference on Coal Llquefactlon, Paio Alto, Ca, May 11-13, 1983. Naiitham, R. V. Ph.D. Dlssertation. Auburn University, AL, 1983. Phiiilp, W. Commun. ACM 1050, 2(12), 12. Sano, Y.; Yamaguchi, N.; Adalchl, T. J . Chem. Eng. Jpn. 1874, 7. 255. Singh, C. P. P.; Carr, N. L. Ind. Eflg. Chem. Process D e s . Dev. 1083, 22, 104. Sapre, A. V.; Gates, B. C. Ind. Eng. Chem. Process Dev. 19818 20, 69. Satterfield, C. N.; Way, P. E. A I C M J . 1072, 18, 305. Shridharani, K. G. Ph.D. Dlssertation, Auburn University, Auburn University, AL. 1983. Tajbl, D. 0.;Slmons, J 6.; Carberry, J. J. Ind. Eng. Chem. Fundsm. 1088, 5 , 171. Wilson, 0. M.; Johnston, R. H.; Huang, S. C.; Tsonopoulos, C. Ind. Eng. Chem. Process Des. D e v . 1081, 2 9 , 94.
Receiued for review September 19, 1983 Accepted September 14, 1984 The authors are grateful to the U.S.Department of Energy for support of this work under Contract No. DEFG2280PC30209.
