1760
Ind. Eng. Chem. Res. 1988,27, 1760-1767
Performance of a Coaxial Gas-Solid Two-Impinging-Streams(TIS) Reactor: Hydrodynamics, Residence Time Distribution, and Drying Heat Transfer Yaniv Kitron and Abraham Tamir* Department of Chemical Engineering, Ben Gurion University of the Negev, Beer Sheva, Israel
The performance of a coaxial two-impinging-streams (TIS) reactor was investigated with respect to the hydrodynamics, the holdup of the particles, their mean residence time, and the residence time distribution as well as the drying heat-transfer process of solids. The results obtained are very useful for practical applications and provide additional insight about the behavior of the reactor. It is concluded that the effective volume for the transfer process is not the actual volume of the reactor but is a certain volume located between the faces of the inlet pipes. Thus, it may be possible to construct very small and effective reactors where their size limit is dictated mainly by hydrodynamic considerations. The coaxial two-impinging-streams (TIS) reactor was first proposed by Elperin (1961) in the early (1960s) as a device for intensifying transfer processes. The essence of the method of impinging streams lies in directing two streams of a suspension of particles one against the other, on the same axis. At the impingement zone of the streams, particles from one stream penetrate into the other stream. This phenomenon yields the following effects which intensify the heat and mass transfer in particulate systems. (a) The relative velocity between the particles and the air is increased due to the countercurrent flow. Under extreme conditions, the relative velocity might reach a value of twice the velocity of the air stream. Thus, the increase of the relative velocity decreases the resistance to the transfer processes, and hence, the process is intensified compared with situations where this effect is absent. (b) The particles oscillate between the faces of the two streams, and hence their mean residence time is increased. (c) The impingement of the streams also causes intensive mixing of the continuous phase (air) which contributes as well to the enhancement of the heat and mass transfer. Among the unique characteristics of impinging streams reactors, the relatively short mean residence time of the solid particles in the reactor, which varies between 0.5 and 2 seconds, is noteworthy. In comparison with the above magnitudes, longer mean residence times of particles in the range 5-65 s were observed in a continuous spouted bed (Mathur and Epstein (1974), p 167) or times of approximately 100 s (Zabrodsky (1963), p 126) and even 500 s were observed in a pilot-plant fludized bed reactor (Kunii and Levenspiel (1969), p 334). Improvements in impinging streams reactors and further research in this field were initiated at the beginning of the 1980s by Elperin and Tamir (1985), who suggested a new configuration of the reactor in which the inlet streams entered the reactor tangentially, rather than a coaxial arrangement. The tangential configuration was explored with respect to its hydrodynamics and residence time distribution (Luzzatto et al., 1984; Tamir et al., 1985), drying of solids (Tamir et al., 1984), mixing of solids (Tamir and Luzzatto, 1985a,b), scale-up (Tamir and Shalmon, 1988), and dissolution of solids (Tamir and Grinholtz, 1987; Tamir and Falk, 1988), as well as combustion (Luzzatto, 1987). In addition, four-impingingstreams and multistage TIS reactors (Kitron et al., 1987) and recently two-impinging-streamsreactors with primary
* To whom
correspondence should be addressed. 0888-5885/88/ 2627-1760$01.50/0
and secondary air feeds (Bar, 1988) were also thoroughly investigated. In order to minimize contact of the phases with the walls of the reactor, the coaxial configuration of impinging jets, rather than the tangential one, was recently applied in absorption and desorption processes (Tamir and Herskowitz, 1985; Tamir, 1986; Herskowitz et al., 1987, 1988), and in the production of emulsions (Tamir and Sobhi, 1985), as well as in combustion of coal (Ziv et al., 1988). The developments in the field of impinging streams until 1972 were listed in Elperin’s monograph (1972), and the more recent ones were summarized by Tamir and Kitron (1987). The major aim of the present research is to investigate more thoroughly the coaxial configuration in order to obtain more insight into the behavior of the reactor with a solid-gas suspension. More specifically, the effect of the distance between the inlet pipes and the mode of introduction of the suspension (cycloned or noncycloned) on the hydrodynamics, the mean residence time of the particles, and their residence time distribution will be studied. In addition, the effect of the volume of the reactor on the drying heat transfer as well as on the above-mentioned design parameters will be explored. This information is very important for practical applications of the TIS reactor, but it is not yet available.
The Reactor The TIS reactor is shown schematically in Figure 1with detailed information about all geometric dimensions. The inlet pipes B are on the same axis and can be moved in order to change the distance between their faces inside the reactor. Two modes of introduction of the solid-air suspension into the reactor are possible. The left-hand side in Figure 1shows the cylone mode, and the right-hand side is the noncycloned feed. These configurations are presented simultaneously for the convenience of the reader, and it is self-evident that both sides were identical when the reactor was tested. As noted, restrictions J were made near the air inlet. This was done in order to create a low-pressure zone at the particles entrance due to the formation of a vena contracta in the air pipe for the purpose of preventing back-pressure effects in the particles’ feed pipe which might cause a nonconstant input mass flow rate. The latter was determined by weighing the amount of particles collected during a certain time interval. The flow rate of the particles was regulated by a rotating ‘‘star’’ feeder mounted at the bottom of a hopper. The air flow 0 1988 American Chemical Society
Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988 1761
R
A l l dimensions i n m m
R1
SOLIDS
Figure 1. Two-impinging-streams reactor. -20
! 41s
Table I. Range of Operating Conditions and Results V, including inlet pipes, m3 (0.65-2.23) X 1, m 0.012-0.112 i03d,, m 1.54, 1.9 Pp, kg/m3 1179, 1153 Was kg/s (8.7-15.8) X Q,m3/s (0-23) x 10-3 kg/s 0-0.037
w,,
P
0-9
V , kg
0.001-0.03
7,
K'
1 &*;d@
lor
0
0
"
~
**' " " 10
ax
M
'
'
V
'
"
mM0
A
~
'
'
1 SOKC
j
'
0
'
WWQ
'
'
5
'
'
s
in cycloned flow at the inlet pipes in noncycloned flow
1
5OOW
Re = DUo/Uo
Figure 2. Euler number versus Reynolds number for air flows at different distances ( 1 ) between inlet pipes for V, = 2.2 X m3.
rate was measured by means of a rotameter. The larger reactor with a volume of 2.2 X m3 (including inlet pipes) is designated by R. In order to test the effect of the volume of the reactor on the hydrodynamics and the heat transfer, its volume was reduced to 0.65 X m3. The smaller reactor R1 is designated in Figure 1 by dotted lines. The attention of the reader is brought to Table I which shows the range of the operating conditions attained in the reactor.
Hydrodynamics The hydrodynamic experiments consisted of measuring the pressure drop on the reactor against the flow rates of air only and the air-particles suspension. The properties of the particles are given in Table I. The pressure drop (aP)was measured between points A1 and A2 in Figure 1 (which were connected together in order to obtain a mean value) and the atmosphere. The measurements were expressed in a dimensionless form of the Euler number defined by Eu = AF'/p,U: (1) and the Reynolds number is given by Re = DU,/v, (2) Figure 2 demonstrates the effect of the following parameters on the hydrodynamic behavior of the reactor for
ua, m/s
Hydrodynamic Experiments
Re Eua (cycloned flow) Eu, (cycloned flow) Eua (noncycloned flow) Eu, (noncycloned flow) AF',mmHzO cycloned flow with and without particle flow noncycloned flow without particle flow noncycloned flow with particle flow
ua,m/s
0.8-1 0.65-0.8 0-36 5000-45000 10-38 2-38 0.9-2 1-5 0-0.36 0-0.157 0-0.36
Heat-Transfer Experiments
Tdi, "C Tde, "c T,, "C T,, "C moisture in millet seeds, kg of water/kg of dry solid, % before drying after drying difference between inlet and outlet
0-20 50-71.1 33.6-57.9 22-27 22-28
28-50 22-46 0-7
Results 9 , J/s
16-3000
h, W/(m2.K) qh, (kg of water evaporated/s)/(mS.K) energy required to transfer the suspension through the reactor, kJ/kg of solids
0-0.06 0.5-2.15
100-1800
air flow: (a) the distance 1 between the faces of the inlet pipes to the reactor, (b) the configuration of cycloned and noncycloned feed of the suspension in the inlet pipes, and (c) the presence of relatively long inlet pipes to provide
1762 Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988
the particles the length needed to reach the velocity of the air stream. The following conclusions may be drawn from Figure 2: (a) The inlet pipes consume a significant portion of the mechanical energy needed to transfer the air through the reactor. For the noncycloned flow with 1 = 0.012 m, it varies in the range 20-33%, while for 1 = 0.112 m the range is 40-57%. For the cycloned flow with 1 = 0.012 m, the range is 12-80%, while for 1 = 0.112 m the variation is between 30% and 90%. (b) The effect of 1 on AP is much more pronounced in the noncycloned flow where a decrease in 1 increases the pressure drop. (c) Values of Eu (or AP) are higher by a factor of 25 compared to a noncycloned flow of the air. Thus, it was impossible to perform experiments at Re > 11000 in the cycloned mode. The effect of 1 on Eu is attributed to the back-pressures exerted on each jet at the impingement zone, which increases the resistance to air flow, and hence the pressure drop on the reactor is increased. It should be noted that the behavior in Figure 2, for the noncycloned air flow, is similar to that obtained in previous investigations (Luzzatto et al., 1984; Kitron et al., 1987; Tamir and Kitron, 1987). The relation Eu-Re consists of two main regimes, the laminar regime and the turbulent one. The laminar regime is characterized by a decrease in Eu versus Re, while the turbulent zone is characterized by a constant Eu and then an increase in its value due to a complicated flow in the reactor. Similarly, in the cycloned mode, the flow is very complicated and thus Eu increases versus Re. The effect of the reactor volume on Eu was also investigated by deto 0.65 X lo9 creasing the reactor’s volume from 2.2 X m3 (R and R1 in Figure 1))namely by a factor of 3.4. It was observed that the increase in the pressure drop on the reactor was approximately lo%, which indicates that the effect of the inlet on the pressure drop is the dominant one. The effect of adding the particles into the air stream is introduced into the hydrodynamicanalysis by applying the following ratios which resulted from a nondimensional analysis (Tamir & Kitron, 1987; Tamir and Shalmon, 1988): (3) 7 = Eu,/Eu, = AP,/AP, CL = Wp/Wa
(4)
The particles used in the experiments were solid grains m (millet seeds) having a mean diameter d, = 1.9 X and density pp = 1153 kg/m3 which were determined as described elsewhere (Kitron, 1987). The 7-p relationship, for cycloned and noncycloned feeds of the suspension, is shown in Figure 3. It covers a wide range of air and particle flow rates, two reactor volumes, and several distances between the faces of the inlet pipes. As observed, there is a conspicuous difference between the behavior of the noncycloned feed and the cycloned feed. In the latter case, the introduction of particles into the air stream reduces 7 and hence the pressure drop. The reduction is sharp for small p’s, and it stabilizes toward a constant value by increasing W,. This behavior may be explained as follows: by increasing the air velocity in the absence of particle flow, turbulence in the air stream is enhanced with possible formation of vortex motion. The introduction of solid particles into the air stream causes part of the energy to be converted into kinetic energy of the particles rather than into expansive turbulence. Consequently, the pressure drop in the presence of particle flow might be lower than in the absence of particle flow. For noncycloned feed,
J d
22-
I
2-
vrX1o4 65
769 1315 17 1
65 65
112 112 11 2
65 1315 22 17 I 22 7 6 9 2 2
112 112 112 112
1315
+ e
CYCLONED FEED
1 ~ 1 0 2
tml
im31
47
22
72
22
52 72
X
536 5 38 605
0
4 7
*
22 22 22
-
04 -
72
112
CYCLONED F E E D
I
I
I I 0
1
2
3
4
1
,
,
5
6
7
,
8
9
1 = wp / w, Figure 3. Euler number ratio versus mass flow rate at various operating conditions for cycloned and noncycloned feed.
the 7-p relationship is linear and the following equation holds: 7 = 0.33~+ 1 (5) As seen in Figure 3, the above equation covers two different volumes of reactors, several distances between the faces of the inlet pipes, and the entire range of particles and air flow rates studied. It should be noted also that Tamir and Shalmon (1988) obtained a single q-p curve in the hydrodynamic experiments in two reactors with geometric dimensions that differed by a factor of 2. Thus, the significance of the q-p lines in Figure 3 is their possible application to reactors with geometric similarity to the reactor in Figure 1. In other words, the prediction of the pressure drop on a reactor in the presence of particle flow can be made from experiments in this reactor in the absence of particle flow as follows: for a given value of 1,the value of 7 = APJAP, is obtained from Figure 3. Experiments conducted with air only yield Eu, = f(Re)and hence AP,for an identical value of Re as desired in the presence of particle flow. An estimation of the errors associated with the hydrodynamic measurements is as follows: The basic quantities, particles, and air flow rates are determined within an accuracy of the order of 2.5% from the mean value, while the accuracy in AP is *5 mmHPO. Thus, orders of magnitudes of average deviations from mean values are *5% for Re, f 1 0 % for Eu, *5% for p, and f15% for 1.
Mean Residence Time of the Particles and Their Residence Time Distributions (RTDs) The mean residence time (7)of the particles was determined by measuring their holdup (V, kg) and mass flow rate (W,, kg/s) and applying the formula 7
=
v/w,
(6)
The holdup was measured for the part of the system including the inlet pipes (B) and the reactor (R or Rl),
1
CURVE 1 I N L E T STREAMS
I
CYCLONED
I
I-
+ x
v.=
538 470
470
B 470
2 2" 1 0 ~ ~ 3 CURVE 2
[: T:!
'12 32 32
112 32 12
'1.2
t L
0 2 1
I'""'
V , = 6 5 ~ 1 0 ~ mS ~T R E A M S
NOT CYCLONED
I'
A
45 03 88
r
~ ~ ~ 1w,,rio3 0 ' lkg/aI Ikglrl
L x io2 Iml
832
1052
112
2 16
102
11 2
I
.".
INLET S T R E A M S ' V CYCLONE0
I
I
06+
04+
/.
28-
A 'Os 0 605
08t
0
37 22
r p = wp / wo
Figure 4. Mean residence time of particles versus mass flow rates ratio at different operating conditions.
namely, between points H and E in Figure 1 for the noncycloned feed and between points G and E for the cycloned feed. The system was brought to a steady-state flow, and at a certain time the particles were collected at the exit (E), while simultaneously their feed into the reactor was stopped. In the case of noncycloned feed, this was done by pressing the flexible pipes (H, Figure 1). In the case of cycloned feed, two metallic cylinders were suddenly introduced at point G where the solid feed entered the reactor. In both cases, the air stream was continued for a few seconds in order to sweep the particles into a collecting bag placed at the exit of the reactor. It should be noted that the hold-up experiments are, relatively, the less accurate ones due to the synchronizationrequired between stopping the feed and collecting the particles. Therefore, 5-10 repetitions were made with a deviation of 50-80% from the mean values at low particle flow rates (-0.002 kg/s) and deviations of 20-30% at larger than 0.01 kg/s particle flow rates. Results for the mean residence time of the particles against p are shown in Figure 4 for various operating conditions. It can be seen that the cycloned feed yielded higher values of 7 by about 0.2 s (-25%). The increase in 7 (or V ) for cycloned feed is mainly due to the spiral trajectory of the particles in the inlet pipes. There are other general conclusions that may be drawn from the determinations of particle holdup and the mean residence time. (a) The holdup of the particles (at constant air flow rate) increases with an increase of W,. At a certain value of W,,, the holdup might reach a maximal value and then even decrease. The latter is due to the increasing influence of inter-particle collisions, and as a result, particles lose their kinetic energy and leave the reactor. This explains also the slight reduction in 7 by increasing p (or W, at constant Wa), as observed in Figure 4. It is also seen that, when air flow rate (W,) is increased, the value of 7 increases. This is due to the increase in kinetic energy of the particles, which enhances their oscillatory behavior and hence causes an increase in 7 . (b) Neither the distance between the faces of the inlet pipes, 1, nor the volume of the reactor, V , (R and R1 in Figure l), affects 7 . This phenomenon suggests that practically the processes in impinging streams are restricted to the zone between the faces of the inlet pipes
1
05
e
15
2
= t / t
Figure 5. RTD curve (V, = 22
X
lo-'
m3).
and also inside the inlet pipes and that, in this zone, intensification of the transfer processes takes place, as observed. Consequentlythe actual volume of the reactor can be reduced at the expense of increasing the pressure drop on it. (c) As seen in Figure 4, the mean residence time of the particles varies between 0.7 and 1 s, while that of the air ( 7 = V,/Q)is approximately 0.03-0.1 s. This difference indicates indirectly that the relative velocity between the particles and the stream of air, hence the intensification in the drying heat-transfer rates later discussed, may be significant. Residence time distribution experiments were performed by using the apparatus developed by Luzzatto et al. (1984). Two syringes were charged with approximately 200 particles taken from the original seeds and colored so as to be easily recognizable. The system was brought to steady state, and then the particles were pushed downward by pneumatic means into the air-particle suspension (at points B1 and B2 in Figure 1)in the form of pulse input. An injection time of 0.08 s was measured by filming the downward motion of the lever and, at the same time an electronic stopwatch, by means of a high-speed camera. It should be noted that the above time is approximately 10% of the mean residence time of the particles. Hence, the pneumatic introduction of the particles can, indeed, be considered a pulse. The colored particles were collected in the sections of a circular plate placed under the solid discharge opening E. I t was possible to determine the RTD curve from the number of the particles in each section, as well as the speed of rotation of the disk. The error in the RTD experiments arises from the relatively small number of tracer particles needed to obtain an ideal pulse. Three to four repetitions were made in each case and the mean distribution was calculated. Thus, the error in estimating 7 was &20%. The residence time distribution curve was described by the following quantities: 9! = t / 7 (7) E(0) = c T / A t (8) The function E(@is defined so that E@)de is the particle fraction at the exit of the reactor which resided in the reactor between the times 0 and e + dt? Figure 5 gives E@) versus 19,which covers all the operating conditions specified in the figure either for cycloned or for noncycloned flow at the inlet pipes. Any attempt to fit the data by stochastic models proposed by Luzzatto et al. (1984) and Tamir and Kitron (1987) failed because of the relatively short values
1764 Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988
of 7 as well as the small magnitude of the variance of the experimental RTD curve. A detailed discussion of this problem is given elsewhere (Kitron, 1987). Therefore, the one-parameter tanks-in-series model (Levenspiel (1972), p 291) widely used to represent nonideal flow was applied. The one parameter of this model is the number of tanks N in this chain, and the exit age distribution function reads
E ( @=
N(Ne)N-l e-NO ( N - l)!
t
,
k
x
X
.
+
(9)
It was found that the best fit of the data is given by N = 17, namely, a series of 17 perfectly mixed tanks. The best fit is shown in Figure 5 where the function E ( @is plotted versus 0. As seen, the distribution curve represents approximately the plug flow behavior of the suspension of the particles in air, but it also contains elements of perfect mixing. Thus, the physical behavior of the TIS reactor may be described as follows: In pipes B (Figure l),the particles are in plug flow so that each pipe may be considered as a plug flow reactor. The impingement zone is the region where perfect mixing takes place. It is located between the two inlet pipes where mixing is caused by the oscillatory behavior of the particles as well as by the two air streams.
Drying of Solids The technique of measuring drymg rates in the insulated TIS reactor was established in the past and is detailed by Tamir et al. (1984) and Kitron (1987). The following points are noteworthy. The experiments were conducted with noncycloned feed in the inlet pipes only and in the constant drying rate regime where the weight of moisturelweight of dry solids, X, varied between 20% and 50%. In addition, the maximal change in moisture content was 7 %. Thus, the surface temperature of the particles, needed for the determination of h or qh, was fixed at the wet-bulb temperature. The latter temperature was measured at point D in Figure l in the presence and also in the absence of particle flow. Dry-bulb temperatures were measured at points F2, Al, A2, C1, C2, and D. Other measurements consisted of the inlet and outlet moisture contents and air and particle flow rates, as well as their holdup. This was done in each run because of the dependence of the above parameters on the initial moisture. Two kinds of millet seeds were used, with d, = 1.9 X and 1.54 X m and pp = 1153 and 1179 kg/m3, respectively. In addition, the drying experiments were carried out for two reactor volumes, namely, VI = 2.2 x m3. The major error in the experiand 0.65 X mental results is caused by slight variations in the air and particle flow rates. Therefore, errors in temperature determinations are within 1-2 K, in particles flow rates are 2-3%, and in the holdup of the particles are 10-20%. The kinetics of the drying heat transfer were determined from measurements of the heat-transfer coefficient defined by Wp(Xi - Xe) h=-- 4 (10) AATlm AATIm where
where i corresponds to the inlet of the reactor, namely, point F2 in Figure 1,and e corresponds to the exit, namely, point D.
154
0
8 7 8 7 158 8 7 158
n
I58
t
8 7
I54 I54
x
A
I
l
l
0
l
l
l
I
l
19 19
19
l
l
20
10
I2 72
22 22 22 22 22 65 65
154
I
l
72
12 32
12 I12
I
l
I , l
~
30
Wp x IO3 (kg/s)
Figure 6. Heat-transfer coefficient versus solids flow rate at various operating conditions.
The surface area of the particles, A , is determined from hold-up measurements and is calculated from A = - 6V dPPP
It is also assumed that A is the actual area because of the relatively small volumetric holdup of the particles in the reactor which is on the order of 2% of the volume of the reactor. The significance of h is 3-fold: (a) From desired operating conditions of q and temperatures, it is possible to obtain from eq 10 the surface area of the particles. If d, is known, it is also possible to obtain the number of the particles and hence their volume. However, note that h does not say anything about the volume of the equipment needed for drying the particles, which is the most important quantity needed for engineering design. (b) h is a key quantity for evaluating hydrodynamic effects associated with a certain flow configuration (such as impinging streams in the present study) and their influence on the heat transfer. (c) It is usually the known quantity available for evaluating the capacity of equipment or for comparing different types of equipment for carrying out a certain heat-transfer process. For practical purposes, the most important quantity is the reactor volume (VI). Hence, it is customary in chemical engineering to define the volumetric heat-transfer coefficient based on VI as follows: 17h
=
E q/x
(
kg of water evaporated/s m3.K
)
(13)
Figures 6 and 7 show the relation between the heattransfer coefficient (h)and the particle mass flow rate (Figure 6) and the holdup (Figure 7), at various operating conditions. It is observed that all the data in the figures point to the same general conclusions: (a) The heattransfer coefficient increases with an increase in the solid particle flow rate and holdup. Eventually, h reaches a constant value at a high W,, where under this condition, the holdup of the particles, V, also attains a constant value.
l
Ind. Eng. Chem. Res., Vol. 27,No. 10, 1988 1765
W0x103 I~Q/SI
-
0.06-
Y
a7
m
E
e
-
'D
2
E -y"
158 8 7 8 7 a 7 158
8 7 158
dpr103 V, ,IO4 Iml 1m31
154 154 154 154 154 19 19 I9
65 65
65 e2 22 22 22 22
I I 10' Im 1
12 112 72
1 2 72
72 I 2 32
0.04 -
a,
X
A 0
*
0
i
v
dpit03 (mi
V, ,lo4 im3)
154 I54 I9 19 I9 19 I54
22
12
22
72 72
15 8 8 7 15 8 8 7 158 8 7 87
154 154
L
22 22 22 22
lo2 Im)
I 2
3 2 72 65 12 6 5 1 1 2 65 72
r
c
002
I 1
A 30
0
10
20
30
10
0
V/Vr ( K g / m 3 )
V x IO3 ( K g )
Figure 7. Heat-transfer coefficient versus solids holdup at various operating conditions.
(b) Within the experimental accuracy, none of the following parameters affect the values of the heat-transfer coefficients: air flow rate which included a 2-fold change in its magnitude; the volume of the reactor which was varied by a factor of 3;and a 10-fold change in the distance between the faces of the inlet pipes. However, the 30% increase in the particle size does not allow one to reach any definite conclusions, despite the fact that, as demonstrated in the figures, the effect of the particle size was negligible. (c) The effective volume for heat- and mass-transfer processes is smaller than the actual volume of the reactor. It is a certain volume between the faces of the inlet and also part of the inlet pipes. Moreover, it is possible to construct very small reactors which are extremely efficient. However, the smaller the size of the reactor, the larger the energy required to transfer the suspension through the reactor for achieving the desired drying load. The scatter of the data in Figures 6 and 7 needs some clarification. The experimental error in the drying experiments is caused by slight instabilities of hot air temperatures, by hold-up measurements with deviations of 20-80% from the mean value depending on the flow rate of the particles, and by technical difficulties in attaining ideal steady-state conditions. The above errors were present at all operating conditions. However, at high particle flow rates, it was more difficult to carry out measurements at steady state than at lower particle flow rates. The dry-bulb temperature gradient on the system was greater than that at the lower particle flow rate by a factor of 10-20 "C. This means that the time needed to arrive at steady state increased as the particle flow rate increased. Thus, if the operating conditions require a long time to arrive at steady state, various parameters might vary, and measurements taken would contain greater experimental error. Bearing in mind the above considerations, we may explain the considerably smaller deviations in h at smaller values of W , or V. In addition, when the overall trends, rather than the point by point results, are considered, it is apparent that the solid curves in Figures 6 and 7 clearly depict the average dependence between the heat-transfer coefficient and the particle mass flow rate and holdup.
20
Figure 8. Volumetric heat-transfer coefficient versus ratio of solids holdup to reactor volume at various operating conditions.
Figure 8 demonstrates the dependence of the volumetric heat-transfer coefficient (qh) on the holdup at various operating conditions. Again, the general trends mentioned before are also repeated here with respect to qh, which is given by qh
= (1.64x io-3)(v/vr)
(14)
Actual values of qh for various drying equipment, such as fluidized bed, spouted bed, etc. (Tamir et al., 1984),indicate that q h varies between 7 X and 2.9 X (kg of water evaporated/s)/(m3-K). In the present case, the maximal value of fh is 0.06,indicating that the TIS reactor may be considered an extremely effective device, on a volumetric basis, for the drying of solids. Figure 8 may also be used for predicting q h for a system with larger volumes, when the holdup of the particles in the new system as well as its volume is known. It is useful to compare the results of the present work with previous investigations. The unequivocal trend of the dependence of the heat-transfer coefficient on the particle flow rate as observed in Figures 6 and 7 is rather new. In the case of Tamir et al. (1984),the number of data points were too limited to draw any conclusion therefrom. However, in the work of Kitron et al. (1987)there is an indication of such a trend. This can be explained by the similarity of the two reactors. In the case of a coaxial TIS reactor, there is one active zone, namely, the impingement zone. In the four-impinging-streamsreactor (Kitron et al., 1987), there are two coaxial pairs (namely, two active zones) which behave similarly to the TIS reactor and hence yield the same behavior of the dependence of h on W,. On the other hand, Meltser and Pisarik (1982)found that for lo3@< 0.825 ( p is the volume fraction of the particles in the feed) h is independent of the particle holdup and that it decreases with increases of the holdup for 0.825 < 103p < 3.5. In the present experiments, the maximal value of is such that the present data are in the same region where the results of Meltser and Pisarik (1982) are correlated by Nu = 0.137Re1.O3
(15)
1766 Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988
where Nu = h d p / k and R e = d,Ua/v,. Thus, the heattransfer coefficient increases with an increase in the air velocity. A plausible explanation for the increase of h with V assumes that, if the holdup ( V ) of the particles is increased, the free space between them is decreased. This is because it has been shown indirectly that the effective volume for the transfer processes where the particles are concentrated is restricted to the volume between the faces of the inlet pipes. In addition, it was shown (Elperin, 1972) that there is a significant increase in the particle concentration of the impingement zone. Thus, an increase of the holdup of the particles in the effective volume causes a decrease in the free volume for the air flow, and hence its velocity is increased. The above result by Meltser and Pisarik (1982), namely, that h U,, gives support to the observation in the present work that h CT V. Finally, it is noteworthy that the order of magnitude of the values of h obtained by Meltzer and Pisarik (1982) is similar to the results of the present work.
Conclusions and Recommendations The performance of the coaxial TIS reactor was explored with respect to the following operating parameters: hydrodynamics, RTD of the particles, and drying heat transfer. The range of the operating conditions which were explored and the results obtained are summarized in Table I, which also gives the reader an idea about the operational capability of the reactor. The most important results obtained were as follows: (1) the holdup of the particles, with air flow rate constant, increases with an increase of particle flow rate. The holdup reaches a maximal value, and then it may decrease. (2) Neither the distance between the faces of the inlet pipes (1) nor the volume of the reactors (R and R1 in Figure 1) has any effect on the mean residence time of the particles. (3) The RTD distribution curve-which was explored over a wide range of operating conditions-approximates a plug flow behavior of the suspension with elements of perfect mixing. (4) It was clearly observed that the heat-transfer coefficient increases with increased particle flow rate or particle holdup. It was also found that the heat-transfer coefficient is independent of the volume of the reactor as well as of the distance between the inlet pipes. In addition, the volumetric heat-transfer coefficients, measured in the present work, were the highest values in comparison with the values obtained in the past, either for TIS reactors or other equipment for drying of solids. The above conclusions led us to describe the behavior of the TIS reactor as follows. The effective volume, in which the heat- and mass-transfer processes take place, is not the actual volume of the reactor. The effective volume is a certain volume located between the faces of the inlet pipes to the reactor and also in some part of them. Hence, it may be concluded that it is possible to construct very small reactors. The size of the reactor is determined from hydrodynamic considerations, namely, the energy needed to transfer the suspension through the reactor. The results obtained from the laboratory reactor may be also used to make estimates for a large-scale reactor. For example, Figure 3 may be used to calculate the pressure drop in a reactor with geometric similitude to the laboratory reactor which operates under identical hydrodynamic conditions as follows: for any desired value of p, it is possible to obtain 11 = APp/APa. AP is determined by studying the hydrodynamics of air onfy in the largescale reactor which yields the value of APa and hence AP,. As seen, Figure 8 describes quite well the results of two different volumes of reactors. Thus, it may be used for
evaluating
qh
if the ratio
v/v,is known.
Nomenclature A = surface area of the particles, eq 12, m2 c = normalized concentration, namely, number of particles collected during time interval At divided by the total number of particles which were collected d, = average diameter of the particles, m D = diameter of the inlet pipe to the reactor, element B in Figure 1, m E(8) = residence time distribution function defined in eq 8 Eu = Euler number defined in eq 1 Eu, = Euler number for air flow only Eu, = Euler number in the presence of particle flow h = heat-transfer coefficient defined in eq 10 1 = distance between the faces of the inlet pipes to the reactor, Figure 1, m N = number of tanks in the tanks-in-series model q = heat-transfer rate, J / s Q = air flow rate at room temperature, m3/s Re = Reynolds number defined in eq 2 U , = air velocity at the inlet pipe to the reactor based on rotameter reading, m/s V = holdup of the particles in the reactor, kg V, = volume of the reactor including inlet pipes, m3 t = time, s Td,, Tdi= dry-bulb temperature at the exit and at the inlet to the reactor, respectively, K T,,, Tk = wet-bulb temperature at the exit and at the inlet to the reactor, respectively, K W,, W, = mass flow rate of the air and of the dry particles, respectively, kg/s Xi, X , = particle humidity (water mass per dry particle mass) at the reactor inlet and outlet, respectively hp = mean pressure drop on the reactor due to friction; it is between points A1 and A2 (Figure 1)connected together and the atmosphere At = time interval between measurements in the RTD experiments ATl,,, = logarithmic mean temperature different defined in eq 11 Greek Symbols pa, pp =
density of air and of particles, respectively, kg/m3
0 = dimensionless time defined in eq 7 = ratio between the mass flow rate of the particles to the
mass flow rate of air, eq 3 7 = ratio between the Euler number in the presence of particle
flow to the Euler number for air flow only, eq 4 volumetric heat-transfer coefficient defined in eq 13 7 = mean residence time of the particles defined in eq 6, s v, = kinematic viscosity of the air, m2/s h = latent heat of vaporization (2545 kJ/kg)
q, =
Subscripts a = air e = exit i = inlet p = particle
Abbreviations
RTD = residence time distribution TIS = two impinging streams
Literature Cited Bar, T. “Hydrodynamics, Heat Transfer and Residence Time Distribution in a Two-ImpingingStreams Reactor with Primary and Secondary Air Feeds”. MSc. Thesis, Chemical Engineering, Ben Gurion University of the Negev, Beer Sheva, Israel, 1988. Elperin, I. T. “Heat and Mass Transfer in Impinging Streams”. Inzh.-Fiz. Zh.1961,6,62-68. Elperin, I. T. “Transport Processes in Impinging Jets”. (in Russian) Nauk.Tehk.Visn. 1972, 1.
I n d . Eng. Chem. Res. 1988,27, 1767-1775 Elperin, I.; Tamir, A. “Method and Reactor for Effecting Interphase Processes”. Israeli Patent 66 162,1985. Herskowitz, D.; Herskowitz, V.; Tamir, A. “Desorption of Acetone in a Two-Impinging-Streams Spray Desorber”. Chem. Eng. Sci. 1987,42,2331-2337. Kitron, Y. “Performance of a Coaxial Gas-Solid Two-ImpingingStreams (TIS) Reactor: Hydrodynamics, Residence Time Distribution and Drying Heat Transfer”. BSc. Project, Chemical Engineering, Ben Gurion University of the Negev, Beer Sheva, Israel, 1987. Kitron, A.; Buchman, R.; Luzzatto, K.; Tamir, A. “Drying and Mixing of Solids and Particles’ RTD in Four-Impinging-Streams and Multistage Two-Impinging-Streams Reactors”. Ind. Eng. Chem. Res. 1987,26, 2454. Kunii, D.; Levenspiel, 0. Fluidization Engineering; Wiley: New York, 1969. Levelspiel, 0. Chemical Reaction Engineering, 2nd ed.; Wiley: New York, 1972. Luzzatto, K. “Investigation of a Two-Impinging-Streams Reactor”. Ph.D. Thesis, Chemical Engineering, Ben Gurion University of the Negev, Beer Sheva, Israel, 1987. Luzzatto, K.; Tamir, A.; Elperin, I. “A New Two-ImpingingStreams Reactor”. AZChE J. 1984,30,600-608. Mathur, K.B.;Epstein, N. Spouted Beds; Academic: New York, 1974. Meltser, V. L.; Pisarik, N. K. “Interphase Heat Transfer Upon Collision between Single and Two-Phase Flows”. Heat Transfer-Sou. Res. 1982,14, 130-133. Tamir, A. “Absorption of Acetone in a Two-Impinging-Streams Absorber”. Chem. Eng. Sci. 1986,41,3023-3030. Tamir, A,; Falk, 0. “Dissolution of Solids and Pressure Drop in Cyclone and Two-Impinging-Streams Semibatch Reactor”. Znd. Eng. Chem. Res. 1988,submitted.
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Tamir, A,; Grinholtz, M. ”Performanceof a Continuous Solid-Liquid Two-Impinging-Streams (TIS) Reactor: Dissolution of Solids, Hydrodynamics, Mean Residence Time, and Holdup of the Particles”. Znd. Eng. Chem. Res. 1987,26,726-731. Tamir, A.; Herskowitz, D. ‘Absorption of COz in a New Two-Impinging-Streams Absorber”. Chem. Eng. Sci. 1985,40,2149-2151. Tamir, A.; Kitron, A. “Applicationof Impinging-Streamsin Chemical Engineering Processes”. Chem. Eng. Commun. 1987,50,241-330. Tamir, A.; Luzzatto, K. “Solid-Solid and Gas-Gas Mixing Properties of a New Two-Impinging-Streams Mixer”. AIChE J. 1985,31, 781-787. Tamir, A.; Luzzatto, K. “Mixing of Solids in Impinging-Streams Reactors”. J. Powder Bulk Solids Technol. 1985,9,15-24. Tamir, A.; Shalmon, B. “Scale-up of Two-Impinging-Streams(TIS) Reactors“. Znd. Eng. Chem. Res. 1988,27,238. Tamir, A,; Sobhi, S. ‘A New Two-Impinging-Streams Emulsifier“. AZChE J . 1985,31,2089-2092. Tamir, A.; Elperin, I.; Luzzatto, K. “Drying in a Two-ImpingingStreams Reactor”. Chem. Eng. Sci. 1984,39,139-146. Tamir, A.; Luzzatto, K.; Artana, D.; Salomon, S. “A Correlation Based on the Physical Properties of the Solid Particles for the Evaluation of the Pressure Drop in the Two-Impinging-Streams Gas-Solid Reactor”. AZChE J. 1985,31,1744-1746. Zabrodsky, S. S. “Hydrodynamics and Heat Transfer in Fluidized Beds”. The M.I.T. Press: Cambridge, MA, 1963. Ziv, A.; Luzzatto, K.; Tamir, A. “Application of Free ImpingingStreams to the Combustion of Gas and Pulverized Coal”. Combust. Sci. Technol. 1988,in press.
Received for review November 24, 1987 Revised manuscript received May 16, 1988 Accepted June 2, 1988
Catalytic Hydroprocessing of SRC-I1 Heavy Distillate Fractions. 7. Kinetics of Hydrogenation, Hydrodesulfurization, and Hydrodeoxygenation of the Neutral Oils Determined by Analysis of Compound Classes and Individual Compounds Sanjeev S. Katti and Bruce C. Gates* Center for Catalytic Science and Technology, Department of Chemical Engineering, University of Delaware, Newark, Delaware 19716
David W. Grandy, Tim Youngless, and Leonidas Petrakis* Gulf Research and Development Company, Pittsburgh, Pennsylvania 15230
The catalytic hydroprocessing of the neutral oils fraction of a heavy liquid prepared from Powhatan No. 5 coal in the SRC-I1 hydroliquefaction process was characterized by conversion in a batch reactor a t 355 “C and 36 atm; the catalyst was sulfided Ni-Mo/y-A1203. Compound classes in the reactant and products were characterized by high- and low-resolution mass spectrometry and lH and 13C NMR spectroscopies. Individual compounds were determined by gas chromatography and gas chromatography/mass spectrometry. The results provide a detailed profile of the hydroprocessing reactions, namely, hydrodesulfurization, hydrogenation, and hydrodeoxygenation. T h e hydrodesulfurization reactions of the major organosulfur components (dibenzothiophene, methyldibenzothiophenes) are characterized by pseudo-first-order rate constants of the order of lo* L/ (g of catalyst-s),and the hydrogenation reactions of major aromatic hydrocarbons (e.g., phenanthrene and pyrene) and hydrodeoxygenation of dibenzofuran and related compounds are characterized by pseudo-first-order rate constants of the order of lo-” L / ( g of catalyst-s). The use of coal-hydroliquefaction products and related heavy liquids as fuels requires their refining by catalytic hydroprocessing to remove sulfur, nitrogen, and oxygen and to increase the hydrogen-to-carbonratios. As part of a systematic investigation of the hydroprocessing reactions of a heavy liquid prepared from Powhatan No. 5 coal by the SRC-I1 hydroliquefaction process, we reported on the separation of 1kg of the liquid into nine chemically distinct
fractions by liquid chromatography (Petrakis et al., 1983a) and a detailed characterization of the fractions (Petrakis et al., 1983b),Katti et al. (1984) reported an investigation of the kinetics of catalytic hydrodesulfurization of the organosulfur compounds in a dilute solution of the neutral oils, the fraction accounting for 72.8 wt % of the coal liquid. The following report is a full description of the catalytic hydroprocessing reactions of this fraction, in-
0 1988 American Chemical Society OS8S-SSS5/88/2627-~~67~Ol.SO~~
