Ind. Eng.
kz = rate constant of reaction between NOz and S032-,L/mol S
NA = absorption rate of SO2 or NOz, mol/cmz s
P = total pressure, atm
p = partial pressure of solute gas, atm QL = liquid flow rate, cm3/s R = gas constant, cm3 atm/mol K S = gas-liquid contact area, cmz T = absolute temperature, K V G = gas flow rate, cm3/s Y = stoichiometric coefficient parentheses ( ) = concentration in gas phase Subscripts
A = gaseous species, SO2 or NO2 B, C = liquid reactant, NaOH and NaZSO3,respectively i = gas-liquid interface 0 = bulk liquid 1,2 = SOz and NOz, respectively L i t e r a t u r e Cited Andrew, S. P. S.,Hanson, D., Chem. Eng. Sci., 14, 105 (1961). Chamber, F. S., Sherwood, T. K., Ind. Eng. Chem., 29, 1415 (1937). Chilton, T. K., Knell, E. W., PACHEC III, 75 (1972).
Chem. Process Des. Dev., Vol.
17, No. 4, 1978
393
Hatta, S.,Tech. Rept. Tohoku Imp. Univ., 8, 1 (1928). Hikita, H., Asai, S., Kagaku Kogaku, 27, 823 (1963). Hikita, H., Asai, S., Tsuji, A,, Preprint of 34th Annual Meeting of the Society of Chemical Engineers, p 64, 1969. Himmelblau, D. M., Chem. Rev., 64, 527 (1964). Kameoka, Y., Pigford, R . L., Ind. Eng. Chem. Fundam., 16, 163 (1977). Onda, K., Sada, E.,Kobayashi, T., Fujme, M., Chem. Eng. Sci., 25, 761 (1970). Onda, K., Kobayashi, T. Fujine, M., Takahashi, M., Chem. Eng. Sci.. 26, 2009 (1971). Ramachandran, P. A,, Sharma, M. M., Trans. Inst. Chem. Eng., 49, 253(1971). Sherwood, T. K., Ind. Eng. Chem., 17, 745 (1925). Sherwood, T. K., Pigford, R. L., Absorption and Extraction", 2nd ed, p 332, McGraw-Hill, New York, N.Y., 1952. Takeuchi, H., Maeda, Y., Itoh, K., KagakuKogaku Ronbunshu, 1, 252 (1975). Takeuchi, H., Ando, M., Kizawa, N., Ind. Eng. Chem. Process Des. Dev., 16, 303 (1977a). Takeuchi, H. Takahashi, K., Kizawa, N., Ind. Eng. Chem. Process Des. Dev., 16, 486 (1977b). Vinograd, J. R., McBain, J. W., J . Am. Chem. SOC., 63, 2008 (1941).
Received for review April 4, 1977 Accepted June 19, 1978
Grateful acknowledgement is made to the Foundation of Development of Technology for Preventing NO,-Emission from Iron Steel Facilities for providing financial assistance for this investigation.
Mass Velocity Effects in an Adiabatic Pilot Plant Simulation of Commercial Catalytic Reforming Antonio E. Eleazar, George W. Roberts, Harold F. Tse, and Robert M. Yarrington' Research and Development Department, Engelhard Industries Division, Engelhard Minerals and Chemicals Corporation, Menlo Park-Edison, New Jersey 08817
Experimental results from an adiabatic, multiple-reactor pilot plant show that both the catalyst aging rate and the reactor temperatures required to produce a given reformate octane number depend strongly on the mass velocity in the reactors. This dependence tended to disappear as the mass velocity was increased above some critical value. The observed effect of mass velocity on catalyst performance is partly attributed to significant differences between the bulk gas temperature and the catalyst surface temperature, especially in the first two reactors of the overall reformer. Data are presented comparing the operation of the adiabatic, multiple-reactor pilot plant with a commercial reformer operating with the same feedstock and catalyst which show that the pilot plant provided an effective simulation of the commercial unit.
Introduction Although mathematical models can play a significant role in the development of new petroleum processes and catalysts (Kmak and Stuckey, 1973), extensive testing in bench scale and/or pilot plant reactors continues to be the backbone of such development efforts. It follows that these experimental reactors must closely simulate commercial plant conditions in order to be reliable devices for process and catalyst studies. Since commercial catalytic reformers for the production of high-octane gasoline or aromatics from naphtha feedstocks are multi-bed reactors, approaching adiabatic operation, a pilot plant of similar design can be an extremely valuable asset in catalyst and process development. Engelhard Industries has operated such multi-reactor, adiabatic pilot plants for several years (Hastings et al., 1972). Despite the advantages of cost and convenience associated with small-scale testing, experimental catalytic reactors can yield misleading results unless care is taken
to understand and minimize effects associated with the low mass (linear) velocities that necessarily exist in small units. These low mass velocities can give rise to significant concentration and temperature differences between the bulk gas stream and the catalyst particle surface. These differences, in turn, can affect the apparent catalyst activity, selectivity, and life. Therefore, in commercial units, where linear velocities are an order of magnitude or more higher than in most experimental reactors, apparent catalyst performance may be significantly different than observed in the laboratory or pilot plant. It is the purpose of this article to demonstrate the effect of interphase heat and mass transfer on the results from a catalytic reforming pilot plant, as well as to show comparisons between the performance of commercial and pilot-plant reactors. The significance of interphase transport in catalytic reforming has recently been discussed by Mahoney (1974). Using pure compounds, he demonstrated a unique, in-
0019-7882/78/1117-0393$01.00/0@ 1978 American Chemical Society
394
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 4, 1978 .-lLe,riI\cs
2
“1hi”Cir.iitX4,Es
ALKYL-
BE\ZIms
I
I
?
Figure 1. General kinetic model of the reforming process.
ternally recycled reactor which can be used to eliminate interphase transport effects. The present paper is based on studies with typical naphtha feedstocks, using conventional tubular reactors. The observations made in this work relate to the effect of intraphase transport on the total series of reforming reactions, as measured by octane number.
Background Catalytic reforming is a process for producing highoctane gasoline, aromatics for petrochemical uses, and hydrogen from a relatively low-octane, low-aromatics hydrocarbon feedstock boiling in the naphtha range (ca. 150 to 420 O F ) . The reforming process plays a key role in refining operations because of its inherent flexibility. All other refining processes (catalytic cracking, alkylation, isomerization, catalytic polymerization, etc.) produce products of relatively fixed octane number. However, reformate octane can be varied as desired, within the constraints of plant hardware. Reforming severity is, therefore, the primary variable available to the refinery operator for adjusting gasoline octane quality. In addition, catalytic reforming of naphtha makes a major part of the hydrogen required for other key operations such as hydrocracking and hydrodesulfurization. Without reformer hydrogen a separate steam reforming plant, consuming large quantities of valuable hydrocarbon feedstock, would be required. The various naphthas which are used as reformer feedstocks are complex mixtures of paraffins, naphthenes, and aromatics, generally containing from six to ten carbon atoms per molecule. In catalytic reforming, the octane number of this feedstock is increased by means of five main reactions listed below and shown in the reaction scheme of Figure 1: (a) dehydrogenation of cyclohexanes to aromatics; (b) dehydroisomerization of alkyl cyclopentanes to aromatics; (c) dehydrocyclization of paraffins to aromatics; (d) isomerization of paraffins to more highly branched paraffins; and (e) hydrocracking of paraffins and naphthenes to increase the aromatics concentration in the reformate. With the exception of reaction (e)-hydrocracking-, all of the above reactions are thermodynamically reversible to some extent under reforming conditions. Detailed understanding of the reaction equilibria is, therefore, very important to successful process engineering. In order to effect all five of these reactions simultaneously, with a single catalyst, that catalyst must have two different types of functionality: (1) a hydrogenationdehyrogenation function, classically provided by platinum or a combination of platinum with other metals, and (2) an acid function, typically provided by the catalyst support, which is usually A1203 or Si02-A1203, to which a halogen is frequently added. Catalysts which contain both
C F
-
CONTROLLER
. FLOW - INDICATOR L . LLVLL P . PRESSURE I
8TAI)ILlZER
REFORMATE
Figure 2. Adiabatic pilot plant flow sheet.
of the above functions are known as “dual-function’’ catalysts. All present reforming catalysts fall into this category. An undesired side reaction in reforming is carbon deposition, which may be represented schematically by heavy hydrocarbons
-
“coke”
+ hydrogen
Coke buildup on the active surface of the catalyst causes the catalyst to lose activity with time on stream. Eventually, regeneration (coke burnoff) is required to restore a practical level of catalyst activity. Coke deposition is favored by high temperatures and by low hydrogen partial pressures.
Mass Velocity Effects A. Experimental Section. 1. Equipment. a. Adiabatic Pilot Plant. The flow sheet for the pilot plant referred to above is shown in Figure 2. Four adiabatic reactors are arranged in series, with heaters located at the inlet of each reactor in order to preheat the feed to the desired inlet temperature. Naphtha is fed from a feed tank through a fixed bed of 13X molecular sieves to remove traces of moisture. The naphtha is then mixed with recycle gas and the mixture is passed through reactors 1 and 2. The pilot plant is equipped so that the recycle gas may be split, a portion mixing with the feed to reactor 1 and the remainder with the effluent from reactor 2. Split recycle is an important feature of the Engelhard/ARCO Magnaforming process. The effluent from reactor 4 passes through a condenser and into a separator. The gas leaving the separator is divided into two streams, a “make gas” and a “recycle gas”. The “make gas” passes to vent through a pressure control valve and a wet test meter. This gas is sampled periodically and the samples are analyzed by gas chromatography. Part of the recycle gas is dried by passing it through a bed of 13X molecular sieve. The remainder is bypassed around the drier to control the moisture content a t the desired level. Although not shown in Figure 2, a duPont Model 510 moisture analyzer can also sample the recycle gas just downstream of the drier. The liquid leaving the separator is continuously distilled to produce a fuel gas, containing mostly C4- components, and a reformate containing mostly C5+ components. The flow rate of fuel gas is measured with a wet test meter and
+ ---
Ind. Eng. Chem. Process Des. Dev., Vol. 17,No. 4, 1978
PREHEAT INSERT
I
Z l
ALUMINA
-CONDENSER
7 A,;
i-----
I.D.
RE001
DRIER
395
1
BYPASS
6 . ~ CONDENSER
MAKE G A S I
U MOISTURE AN A LY Z E R
COMPRESSOR
F E E D PUMP
SE~ARATOR
Figure 4. Flow sheet for isothermal tubular reactor. Table I. Experimental Conditions for Mass Velocity Tests in Adiabatic Pilot Planta sumrficial mass feed velocity, rate, g/h lb/h f t 2
68 190 489
Figure 3. Reactor insert.
samples of this gas are periodically analyzed by gas chromatography. The research octane number of the C5+ reformate is measured by means of a Waukesha CFR octane engine using ASTM procedure D2699-70. The measured octane number is corrected for C4- in the reformate and for C5+ in the fuel gas. The individual reactor baskets (Figure 3) are 11/16 in. i.d. with a 5/32 in. 0.d. thermowell. Each reactor is packed to a total volume of 60 cm3 with a mixture of catalyst and inert diluent. The diluent is 8 to 14 mesh Alcoa T-61 tabular alumina mesh granules. A short section of diluent is packed between the catalyst bed and the metal spacer to prevent axial heat loss from the catalyst bed. Skewed loading of catalyst in the reactors is usually used to conform to commercial procedures. In each reactor, the gas is preheated to the desired inlet temperature by being forced to flow against the tube wall by a helical insert. Both the helix and the basket are inserted into the reactor tube, which is surrounded by a reactor block. The block comprises both thermally conductive and insulating material. The former is bronze, and serves to spread the heat evenly in a given zone to eliminate hot spots. The insulating material, marinite, isolates the various zones. The block is heated by a four-zone furnace. The control thermocouples are inside the block, near the center of each zone. The temperature profile through the bed is initially forced to an isothermal condition during a calibration step included in the start-up procedure. The block temperatures measured during the isothermal calibration period are used to adjust to zero heat loss during adiabatic operations.
50 130 360
weight of catalyst, g
R,
R,
R,
R,
total
2.92 5.84 12.50 1.66 2.08 4.67 5.83 8.16 16.34 35.00 12.00 15.00 21.00 42.00 90.00
a Process conditions: total pressure = 1 8 0 psig inlet; 1 5 0 psig separator; recycle ratio = 10/1 gas/feed; 50% to R , , 50% t o R,; target octane = 99.0 RONC; weight hourly space velocity (WHSV): overall, 4.0 h - ' ; R , , 30 h - ' ; R,, 24-l; R,, 17 h - l ; R,, 8.6 h - I ; reactor temperatures: R , = 880 "F inlet; R , = 910 "F inlet; R, = R, = as required to meet target octane.
The adiabatic pilot plant is computer monitored and controlled by a GE/PAC 4020 process control computer. b. Isothermal Pilot Plant. This is a single reactor unit, as described previously by Hettinger et al. (1955); however, in the present experiments, a recycle loop as added. The inside diameter of the reactor tube is 0.957 in. and the outside diameter of the thermowell is 0.25 in. A flowsheet of the isothermal unit is shown in Figure 4. 2. Catalysts. The catalysts used in the adiabatic and isothermal pilot plant runs described in the following section were bimetallic catalysts, with the physical properties as follows: shape, extrudate; nominal diameter, l/ls in.; nominal length, 1/4 in.; bulk density, 0.645 g/cm3. 3. Experimental Design. a. Adiabatic Pilot Plant. A typical Midcontinent naphtha was used for these runs. Test conditions are listed in Table I. All runs were at 150 psig with 10/1 gas/feed recycle ratio. The recycle was split, as in commercial Magnaforming. For the pilot-plant experiments, 50% of the recycle went to R1 and 50% went to R3. The target octane value of 99.0 RONC was maintained over the length of the run by adjusting the inlet temperature to R3 and R4 as required. Inlet temperatures to R1 and R2 remained constant at 880 and 910 O F , respectively. A series of experiments was performed at a constant overall weight hourly space velocity (WHSV) of 4 h-' (g of naphtha/g of catalyst, h), and with a constant space velocity in each of the individual reactors, as given in Table
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Table 11. Experimental Conditions for Mass Velocity Tests in Isothermal Pilot Plant0
I
superficial mass velocity, lb/h f t 2
feed rate, glh
wt of catalyst, g
25 75 150
50 150 300
8.33 25.0 50.0
(I Process conditions: total pressure, 250 psig; recycle ratio, 3/1 gas/feed; target octane, 100 RONC; weight hourly space velocity, 6.0 h-’; reactor temperature, as required to maintain target octane of 100 RONC.
AGING
T I M E , HR
Figure 5. Pictorial representation of start of run temperature and cycle length.
I. In these experiments, the superficial mass velocity was varied from run to run by simultaneously adjusting the naphtha feed rate and the total catalyst charge in direct proportion. The catalyst charge was varied by changing the catalyst/diluent ratio. b. Isothermal Pilot Plant. A similar series of runs was also made in the isothermal pilot plant, using the same catalyst and a similar feedstock. The reactor temperature was adjusted during the course of each run in order to maintain the target reformate octane number, 100 RONC. Experimental conditions are shown in Table 11. 4. Results. a. Adiabatic Pilot Plant. Three criteria were used to determine the influence of mass velocity on catalyst performance: (1) the initial apparent catalyst activity, ( 2 ) the apparent catalyst aging rate, and (3) the initial yield of C,+ reformate. In the present case, the initial apparent catalyst activity is measured by the inlet temperatures to reactors 3 and 4 which are required to produce 99 RONC reformate at the beginning of the run. This temperature is referred to as the “start-of-run (SOR) temperature”. The apparent catalyst aging rate is expressed in terms of the cycle length, which is defined as the elapsed time between the start of
the run and the point at which the C,+ reformate yield had decreased by 2.0 vol 70from its start-of-run value. Figure 5 illustrates the way in which SOR temperature and the cycle length are derived from the experimental data. Figure 6 shows how the SOR temperature and the relative cycle length depend on mass velocity for the present experiments. As the superficial mass velocity is increased from 68 to 489 lb/h ft2, the SOR temperature decreases by 52 O F , from 1010 to 958 O F . Similarly, this increase in mass velocity causes the cycle length to increase by a factor of 6.3. Note in Figure 6 that the effect of mass velocity on both temperature requirement and cycle length is more pronounced at low mass velocity. The curves in this figure suggest that the influence of mas velocity on catalyst performance will disappear entirely once some critical value of the mass velocity is exceeded. Reformate yield (reaction selectivity) also showed a dependence on mass velocity. For instance, the initial C,+ reformate yield increased from 88.0 w t 70at 68 lb/h ft2 to 89.3 wt % at 489 lb/h ft2. Such yield differences are very important to the refiner and are of great significance in assessing the performance of experimental catalysts. b. Isothermal Pilot Plant. The results of the experiments listed in Table I1 are given in Table 111 and Figure 7. Table I11 shows that the SOR temperature, and therefore the initial apparent catalyst activity, had a very weak dependence on superficial mass velocity in the isothermal reactor runs. The isothermal temperature
1040
-
- 6
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-
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z
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/
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ADIABATIC REACTORS WHSV = 4.0
- 4
z a 3
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a - 2
960-
940
-
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0
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S U P E R F I C I A L M A S S V E L O C I T Y , LBS
400
300 HR
10
I
L
100
-F
T ~
Figure 6. Effect of mass velocity on start-of-run temperature and catalyst aging rate.
500
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 4, 1978 397 3.0
-
T I-
2.0-
W -1 W -I
PRESSURE
*
V V
W
?
I-
4
g
2 5 0 PSlG
RECYCLE RATIO 3/I FEED MIDCONTINENT N A P H T H A WHSV 6 HR-' 1.0
-
0
20
I
I
,
40
60
80
I
100
S U P E R F I C I A L M A S S VELOCITY,
120
LBIHR-
140
I
I60
I 180
I 200
F T ~
Figure 7. Effect of mass velocity on catalyst aging rate in an isotherma! reactor. Table 111. Effect of Mass Velocity on Apparent Initial Catalyst Activity-Isothermal Reactor mass velocity, Ib/h ft2
initial temperature required for 100 RONC, OFU
25 75 150
954 952 948
Average over first 1 2 t o 18 h of run, corrected to 100 RONC.
requirement decreased only 6 O F as the mass velocity was increased sixfold, from 25 to 150 lb/h ft2. This behavior is in strong contrast to that observed in the adiabatic unit, where the temperature requirement was very sensitive to mass velocity. Figure 7 shows that mass velocity once again had a pronounced effect on the apparent catalyst aging rate. The observed catalyst life more than doubled when the mass velocity was increased from 25 to 150 lb/h ft2. 5. Discussion of Results. There are several hypotheses that must be considered in attempting to explain the observed effects of mass velocity in the adiabatic unit. Among these are environmental effects related to differences in chloride retention on the catalysts, as well as mechanical effects due to the type of pump used. These two phenomena were considered unimportant compared to the two transport limitations listed below: (a) temperature and concentration differences between the bulk gas and the geometric catalyst surface, and (b) variations in gas-solid contacting, due to changes in the catalyst/ diluent ratio. Calculations performed using the procedures suggested by Mears (1971) and Satterfield and Sherwood (1963) show that significant temperature gradients did exist between the bulk gas the the geometric catalyst surface, expecially in the first two reactors. These gradients are discussed in detail in the following paragraphs. Unfortunately, for adiabatic reactors, the effect of catalyst dilution is difficult to analyze in a quantitative manner. However, some guidance might be obtained by applying criteria derived for isothermal reactors (Mears, 1971). Such calculations show that catalyst bypassing could have had a significant influence on conversion,
especially at the lowest mass velocity and in the first reactor. Furthermore, the predicted effect of catalyst bypassing is directionally consistent with the trend of the data in Figure 6. That is, according to theory, catalyst bypassing should become less severe at the ratio of diluent to catalyst is decreased. For the particular experimental design shown in Table I, the diluent to catalyst ratio decreases as the mass velocity increases, so that the decreased temperature requirement at high mass velocity that is shown in Figure 6 might be due, at least in part, to decreased catalyst bypassing at the higher mass velocities. A more definitive statement is not possible at this time, since a quantitative theory of bypassing in adiabatic catalyst beds has not yet been developed. The significance of interphase (bulk gas to catalyst particle surface) temperature differences can be analyzed in detail (Mears, 1973), and it can be shown that these temperature differences can have an important effect on catalyst performance. This analysis is the subject of the following section. External Temperature Differences. Temperature differences between the bulk gas and the geometric particle surface were estimated from a steady-state enthalpy balance on the catalyst particles, i.e. AT = T g- T , = q / h a
(An
where q is the enthalpy change due to reaction per unit volume of catalyst, h is the heat-transfer coefficient, and a is the external surface area of the catalyst particles per unit bulk volume. For the present catalysts, a E 320 ft2/ft3 of reactor volume. Of the five main reforming reactions listed previously, the first three-dehydrogenation, dehydroisomerization, and dehydrocyclization-are strongly endothermic. These three reactions usually dominate the overall enthalpy effect, so that the overall reaction is net endothermic. In such a case, the value of q is greater than zero, and the surface temperature of the catalyst particles is less than the bulk gas temperature. The heat consumption rate, q, varies from run to run, from reactor to reactor, and with axial position within any given reactor. Values of q may be calculated from the experimentally measured temperature profiles in the
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Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 4, 1978
Table IV. Heat Consumption Rates and Temperature Gradients (superficial Mass Velocity = 1 9 0 Ib/h ft') heat consumption rate, Btulh f t 3 of catalyst x reactor
R, Rl R3 R4
t o p 'I4 of bed 49 2 245 134 44
Table V. Heat-Transfer Coefficient: Gas to Solid Particles
temperature gradient, OF, (bulk gas to surface of particles)
total bed
top 'I4 of bed
total bed
203 107 71 23
30.8 15.3 8.4 2.8
12.7 6.7 4.4 1.4
feed rate, glh
G mass velocity lblh f t 2
NRea
Nprb
hC
hjd Btulh f t 2 "F
50 140 360
68 190 489
18 49 124
1.8 1.8 1.8
0.62 0.33 0.20
34 50 78
a NRe = d,G/p. Npr = C,p/k. j H was obtained from the graph of Mears (1971, p 33) (Colburnj factor). h = jHCpG/NprZ'3.
g
900h
20
SUPERFICIAL MASS VELOCITY 190 Ib/lt2,hr
2 6 0 H O U R S ON STREAM Y
3 e
t
: IO t 3 REACTOR I
' 0
02
06
04
FRACTIONAL
08 DISTANCE IN BED
Figure 9. Temperature gradients in reactors 1 and 4 between bulk gas and the particle surface.
1 $-
40
I \
\
REACTOR I
I
0
I
200
I
1
400
600
- IbS/hr,ft* Figure 10. Effect of mass velocity on temperature gradients between bulk gas and the particle surface for top quarter of each reactor. SUPERFICIAL
MASS VELOCITY
cP; Prandtl number = 1.8. The particle Reynolds number was calculated using the diameter of a sphere of equivalent surface as the characteristic dimension. The heat-transfer coefficient values calculated for each reactor are listed in Table V. From the temperature profiles of Figure 8, the temperature gradients between the bulk gas and the surface of the particle have been calculated for various bed depths in reactors 1 and 4. These gradients are seen, in Figure 9, to change throughout the bed by an order of magnitude and between beds, becoming insignificant in the last reactor. The effect of mass velocity on these gradients is shown in Figure 10, where the gradients across the top quarter of the bed have been plotted at the three levels of mass velocity. A t the highest mass velocity used, the gradients
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 4, 1978 Table VI. Average Temperature Gradients in Top Section of Commercial Reactors bulk-to-particleotemp mass difference, F velocity, l b / h f t 2 N R ~ jH h R, R, R, R, ~
380 634
1500 2500
0.12 0.086
144 172
6.1 5.1
5.4 4.5
1.5 1.2
0.4 0.3
are approaching insignificant levels after reactor 1. Commercial reactors usually operate a t mass velocities above 1500 lb/h ft2. Table VI lists the temperature differences expected in commercial reactors at two values of mass velocity. At industrial conditions, the catalyst surface temperature should be quite close to the bulk gas temperature, although even the differences shown in Table VI might have some effect on performance. In order to appreciate the effect on catalyst performance of low catalyst surface temperatures caused by low mass velocities, both the kinetics and the thermodynamics of the reactions that occur in the lead reactors must be considered. The temperature dependency for the kinetics of naphthene dehydrogenation reactions that predominate in these front reactors has been presented by Smith (1959). The effect of a 20 O F drop on the reaction rate constant for dehydrogenation is listed below. The rate constant ~
~
temperature,
O
F
reaction rate constant, mol/h lb of cat. atm
900 880
0.096 0.066
is about 30% less a t the lower temperature. These rate constants, in conjunction with the calculated gas-to-catalyst temperature differences, support the hypothesis that low catalyst particle temperatures can have a pronounced effect on apparent catalyst activity. In order to facilitate thermodynamic analysis of the reforming reactions, consider the greatly simplified reaction sequence -HZ
-3H*
P = C P C A +3H2 R
where P represents paraffins, CP represents cycloparaffins, and AR represents aromatics. At low temperatures, such as are characteristic of the first reactor outlet, the aromatics/cycloparaffin equilibrium is shifted toward the cycloparaffins, and the cycloparaffin/paraffin equilibrium is shifted toward the paraffins. In fact, if the catalyst temperature is allowed to become too low, some of the aromatic compounds, especially the lighter ones, may be in chemical equilibrium with their corresponding cycloparaffins. In such a case, a ring-opening reaction, in which cycloparaffins are hydrocracked to paraffins, is the only reaction involving ring compounds which is thermodynamically allowable. In reforming, these ring-opening reactions are highly undesirable because they cause the octane number to decrease, and because the paraffins that are formed cannot be converted back into ring compounds in the downstream reactors without a serious loss of selectivity. In order to compensate for any ring opening that may occur in the lead reactors, the temperatures in the downstream reactors must be increased to favor dehydrocyclization, both kinetically and thermodynamically. However, these higher temperatures lead to an increased rate of hydrocracking, as well as to an increased dehydrocyclization rate. The presence of paraffin and cycloparaffin hydrocracking
399
prevents a one-to-one reconversion of the ring compounds which were lost due to excessively low temperatures in the lead reactors, thereby leading to decreased reformate yields. Thus, the speculation that low mass velocities lead to increased ring opening in the first two reactors because of lower average catalyst temperatures is directionally consistent not only with the observed effect of mass velocity on apparent catalyst activity, but also with the observed effect of mass velocity on C,+ yield. The phenomena underlying the increased rate of catalyst aging at low mass velocities are not quite so clear. Certainly, a portion of this increased aging rate is related to the higher SOR temperatures that were required in the low mass velocity runs. These higher temperatures would tend to increase the rate of coke deposition and the rate of metal crystallite growth. However, it can also be speculated that the increased aging is partly related to the transport of coke precursors which are formed within the catalyst particles. The concentration of these coke precursors at the external pellet surface, and therefore within the pores of the catalyst, should increase as the mass velocity is decreased; this is because the decreased mass velocity reduces the mass-transfer coefficient between the particle surface and the bulk gas. An increased concentration of coke precursors would, of course, result in an increased rate of coke laydown and an increased apparent aging rate. The various differences and similarities in behavior between the adiabatic and isothermal units are not easy to rationalize. Specifically, although yield and SOR temperature are much stronger functions of mass velocity in the adiabatic units than in the isothermal reactors, it is nevertheless observed that mass velocity has a similar effect on the apparent catalyst aging rate in both units. This situation tends to support the speculation that the aging and activity effects are linked to different phenomena. Thus, if the effect of mass velocity on aging is primarily related to transport of coke precursors, it would not be surprising to observe similar aging behavior in the two types of unit. Moreover, if the effect of mass velocity on apparent catalyst activity and selectivity is primarily related to low-temperature ring opening in the lead reactors, then the isothermal reactor, which operated 70 O F and more hotter than the first adiabatic reactor, should be relatively immune from the effects of mass velocity, whereas the adiabatic unit would show a strong sensitivity. The observed differences between the adiabatic and the isothermal reactors serve to emphasize the utility of the adiabatic pilot plant as a tool for meaningful catalyst testing and proces development. These differences also support an old bromide in chemical reaction engineering, to the effect that it is always desirable to study a given reaction in at least two different types of reactor, in order to ensure that all factors which can influence performance have been adequately investigated.
Comparison of Adiabatic Pilot Plant with Commercial Units Pilot plant tests were made with a feedstock obtained from a commercial unit. The process conditions being used in the plant at the time the feedstock was sampled became the conditions for the pilot plant run. The objective of the pilot plant run was twofold. The adiabaticity of a small pilot unit required confirmation. Since the commercial units are adiabatic reactors, comparison of the temperature drops between the commercial and pilot plant units would provide a method of comparing
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Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 4, 1978
Table VII. Operating Conditions for Comparison of Adiabatic Pilot Plant with a Commercial Reformer unit catalyst data cycle age
commercial unit
pilot plant
4 1 start t o mid-cycle start of cycle
WHSV:
total reactor 1 reactor 2 reactor 3 reactor 4 recycle ratio, H,/HC reactors 1 and 2 reactors 3 and 4 inlet temperature, F reactor 1 reactor 2 reactor 3 reactor 4 inlet pressure, psig reactor 1 reactor 2 reactor 3 reactor 4 target octane number (RONC) superficial mass velocity (Ib/h ft2)
1.82 h-' base base base base
1.86 h-I base base base base
4.1 6.4
4.0 5.8
base base base base
base base- 17" base- 1 5 base- 1 5
395 380 370 360 92
384 384 362 362 92
1000-2000
200
Table VIII. Temperature Drop for Commercial Unit and Adiabatic Pilot Plant
reactor 1 2 3 4 total
commercial unit. " F
pilot plant, "F
105 75 27 17 224
109 63 29
8 209
heat losses in the pilot plant. The other objective was to compare yields at the same octane number. The operating conditions used in the pilot plant in relation to plant conditions are listed in Table VII. These were set as close as possible to plant conditions. Pressure declines through a commercial unit due to frictional losses in the heaters and catalyst beds. In the pilot plant, very little pressure drop is observed due to the low mass velocity. An artificial pressure drop is therefore established in the pilot plant by operating reactors 3 and 4 at a lower pressure than for reactor 1 and 2. The temperature in reactor 1was set identical with the commercial unit. The other temperatures were adjusted to obtain target octane. Differences between the plant and pilot plant temperatures reflect the age difference between the catalysts. The temperature drop in the first bed was quite close for the two units (Table VIII). This verified that the pilot plant operation was, in fact, very close to adiabatic. The overall temperature drop was also remarkably similar for the two units, as were the drops for reactors 2 through 4. The yield results are presented in Figure 11,where plant results for about the first half of a cycle are shown in comparison to the steady-state pilot plant results. The C,+ yields are slightly lower for the adiabatic pilot plant run.
0
1
I
0 t
i 2urn i B
P t :*
=
.. _ _ - _.. , _
+loo-
* .
. . . . - -. - - .. . *.
*
-1001
3
4
M O h T H S ON
3
6
STREAM
Figure 11. Yields in a commercial reformer and in the adiabatic pilot plant.
This is due in part to several factors: (a) fresh catalyst was used in the pilot plant, (b) the mass velocity is possibly slightly on the low side in the pilot plant equipment, and (c) errors in material balances in the plant and pilot plant. Even with these factors, the agreement in yields is remarkably close. Conclusions Reliable testing of reforming catalysts, as well as the development and optimization of reforming processes, requires the proper experimental equipment and an appreciation of the artifacts that can be introduced by operation at low mass velocities. The present results have demonstrated that an adiabatic, multi-reactor pilot plant can provide a good simulation of commercial reformers. The strong potential influence of mass velocity on apparent catalyst activity and catalyst aging has also been demonstrated, and it is clear that accurate results can be obtained only when the mass velocity is sufficiently high so that its influence is negligible. Catalyst bypassing, caused by the relatively high diluent to catalyst ratios that were used in the present study, may have been partially responsible for the observed results. Further development of the theory of catalyst contacting in adiabatic reactors is required before this effect can be quantified. Literature Cited Hastings, K. E., Pfefferle, W. C., Tseng, Y. M., Res. Dev., 23 (3),15-19 (1972). Hettinaer, W. P.. Jr., Keith, C. D., Grinq, J. L., Teter, J. W., Ind. Eng. Chem.,
47,719 (1955). Kmak, W. S., Stuckey, A. N., Jr., "Powerforming Process Studies With a Kinetic Simulation Model", Paper 56a, AIChE National Meeting, New Orleans, La., Mar 14. 1973. Mahoney, J. A,, J . Catal., 32, 247-253 (1974). Mears. D. E.. Ind. Eng. Chem. Process D e s . Dev., 10, 541 (1971). Mears, D. E., J . Catai., 30, 1283-1287 (1973). Satterfiild, C. N., Sherwood, T. K., "The Role of Diffusion in Catalysis", pp 29-34, Addison-Wesley Publishing Co., Reading, Mass., 1963, Smith, R. B., Chem. Eng. Prog., 55 (6), 76 (1959).
Received for review April 13, 1977 Accepted April 17, 1978 Presented at the Petroleum Session, First Chemical Congress of the North American Continent, Mexico City, Mexico, December 1975.
