Ind. Eng. Chem. Res. 1992,31, 2265-2272
2265
Modeling the Thermal Cracking of Ethane and Propane in a Non-Isothermal Vertical Pneumatic Transport Reactor Hiroki Koyama and Joshua S. DranofP Department of Chemical Engineering, Northwestern University, Euanston, Illinois 60208-3120
A model has been formulated for noncatalytic thermal cracking of ethane and propane in a nonisothermal, adiabatic, vertical, pneumatic transport reactor. This model takes into consideration hydrodynamic properties of the gas-solid mixture, gas-particle heat transfer, and a kinetic scheme involving 81 radical reactions among 11molecules and 11radicals. Simulations of ethane cracking at a fixed conversion show that the particle diameter and solid to hydrocarbon ratio have the greatest effect on reactor performance; an optimum particle diameter of 50 pm was found. Performance of a simulated heater coil reactor in a conventional pyrolysis furnace was compared with that of the transport reactor. Calculations show that the transport reactor can achieve greater ethylene yield at much shorter residence times for a given conversion. However, propylene, which is one of the main products in propane cracking, is produced at a higher yield in a conventional pyrolysis furnace.
Introduction In a vertical pneumatic transport reactor (VPTR), solid particles, which may be reactants, catalysts, or inert heat carriers, are physically carried upward through a transport line by cocurrently flowing reactant gas. The gas velocity in such cases is much higher than that required for minimum fluidization, and the reactor bed consequently operates in a region of dilute phase transport. Vertical pneumatic transport (riser) reactors have been utilized in the petroleum industry for the catalytic cracking of gas oil and resid to gasoline. The VPTR concept can also be applied to homogeneous chemical reactions such as thermal cracking of hydrocarbons, using inert solid particles supplying the heat requirements of the reactions. conventional pyrolysis furnaces for such service can have only limited operation at high severities because of the problem of reactor coil fouling as well as limitations of heat flux and coil metallurgy. However, a transport reactor system allows for potential continuous operation at severities above those possible in conventional pyrolysis furnaces; the heat of pyrolysis can be provided by the entrained hot inert solids rather than transfer through the reactor coil walls. Coke deposition will still occur, but predominantly on the hot solids rather than the cooler tube walls. The entrained, coke-coated solids leaving the transport reactor can then be conveyed to a regenerator where the coke is burned off and the particles are simultaneously reheated for return to the reaction zone. The Thermal Regenerative Cracker Process, developed by Gulf Oil Chemical Company (Ellis et al., 1982; Gartside and Ellis, 1983) for thermal cracking of heavier hydrocarbons, is an adaptation of this concept. Because thermal cracking of heavier hydrocarbons such as gas oil produces coke easily and causes severe coil fouling, a conventional pyrolysis furnace for such service will need frequent shutdowns for decoking. Most developments of thermal cracking processes have been focused on achieving short residence times in order to yield maximum olefin products. For example, Kellogg developed the “millisecond furnace” (Ennis et al., 1975), an otherwise conventional pyrolysis furnace in which small-diameterstraight coils are used to increase heat flux. However, it requires a large number of passes to compensate for the small capacity of each coil. Other minor improvements have resulted from the use of expanding coils and split coils in the pyrolysis furnace. These reduce
* To whom correspondence should be addressed.
the maximum coil skin temperature because mass velocity decreases along the coil; they also reduce the pressure drop which leads to improved selectivity. However, coil fouling remains a problem. In order to eliminate the limitation of heat flux through the coil, novel pyrolysis reactor technologies have sought to provide internal heat sources by means such as partial oxidation (an autothermal reactor), hot liquid (a molten bed reactor), or hot solids (a fluidized bed reactor) (Ross, 1983). Since the vertical pneumatic transport reactor (VPTR) has the potential to improve olefin selectivities in the pyrolysis of hydrocarbon feedstocks, this work was undertaken to study thermal cracking in such a device by means of computer simulation. The emphasis was on formulation of a model for light hydrocarbon cracking taking into account reaction kinetics, fluid mechanics, and heat transfer in gas-particle systems, and subsequent investigation of model predictions under various conditions. There were several previous models developed to account for the reactant conversion in the VPTR for (1) catalytic reactions (Pratt, 1974; Paraskos et al., 1976; Varghese and Varma, 1979; Fan, 1981; Fan and Hwang, 1981; Shaikh and Carberry, 19841, (2) noncatalytic gassolid reactions (Fan et al., 19841, and (3) noncatalytic gas-phase reactions (Jepson, 1986). Fan (19811, Fan and Hwang (1981), Fan et al. (1984), and Jepson (1986) incorporated empirically obtained hydrodynamic properties in their modeling efforts. However, there have been very few experimental measurements of heat-transfer rate between gases and suspended fine particles, and only limited correlations are available (Bandrowski and Kaczmarzyk, 1978; Kat0 et al., 1983). To date, only Jepson (1986) has developed a non-isothermal transport reactor model incorporating an empirical gas-particle heat-transfer correlation. In the present study, an improved non-isothermal VPTR model was developed for the noncatalytic thermal cracking of ethane and propane in the gas phase. Since only gasphase reactions are considered, the reactor conversion is related to gas temperature, gas pressure, and gas residence time. The model takes into consideration (1)the fluid mechanical properties of the gas-particle system, (2) the initial particle velocity and bed configuration, and (3) the gas-particle heat-transfer process. The effects of various reactor parameters on the olefin product yields at a given reactant gas conversion have been studied using this model. In addition, the product yields predicted by the VPTR model and a kinetically similar model of a conventional pyrolysis furnace have been evaluated and compared.
0888-5885/92/263~-2265~03.o0/0 0 1992 American Chemical Society
2266 Ind. Eng. Chem. Res., Vol. 31, No. 10, 1992
gravity, wall friction, and acceleration:
Table I. Molecular and Radical Components in the Reaction Scheme for Thermal Cracking of Ethane and Propane (after Sundaram and Froment, 1978) components radicals
2fgPgUe2 PA1 - Of,U,2 + --dPt = PA1 - 4g + + dz
H', CHE C,H',', C2H2, CSH6., l-CSHT', 2-C3H7', C4H7', 1-C4Hg', 2-C4Hg', C5Hll'
The Model The reactions considered in this work describe the gas-phase thermal cracking of ethane and propane. The heat required for these endothermic reactions is provided by spherical, hot, inert solid particles carried through a vertical pneumatic transport reactor by cocurrently upward flowing reactant gas. The kinetic model for the thermal cracking reactions developed by Sundaram and Froment (1978) was incorporated in this work. Since they have previously discussed their model in some detail, only the major feature will be described here. The Sundaram-Froment radical reaction scheme for ethane-propane cracking considers 81 radical reactions (neglecting the coke formation reaction) involving the 11molecular and 11radical species listed in Table I. The reactions are all first or second order in concentration. This model represents a successful compromise between the much more comprehensive SPYRO model (Goosens and Dente, 1978; Gooeena et al., 1978) with its 385 reactions and simpler, less accurate models suggested in the literature from time to time. Assuming that the gas and solid phases are in plug flow, the continuity equation for the component j in the multicomponent system can be written as
The rate of the ith reaction is expressed as
Dt
Here fg and f, are the gas-wall and solid-wall friction factors, respectively. Considering the momentum balance for the solid phase alone taking into account the drag force, the solid velocity is given by (Yang, 1977)
where CDs is the drag coefficient for a single spherical particle. The solids are initially accelerated by the drag force exerted by the gas in the reactor entrance region; subsequently the relative velocity between gas and solids approaches the terminal velocity in a multi-particle system given by eq 9 (assuming that the left-hand side of eq 8 is
zero). The acceleration of the solid particles in the entrance region is referred to as the "entrance effect" in this work; the length of the entrance region depends on the initial solid velocity when the other parameters are fixed. In order to study the significance of the entrance effect on reactor performance, minimum and maximum values of the initial solid velocity were considered. The minimum initial solid velocity is obtained by assuming an initial void fraction of 0.45 (which is about the void fraction for minimum fluidization) in the mass balance equation for the solid phase as suggested by Yang (1977): 4
t = l -
(P,
with (3) i=l
The reaction rate constanta ki follow the usual Arrhenius form. aij is the stoichiometric coefficient of species j in the ith reaction, and ujj is the reaction order for species j in the ith reaction. (Note: Both cyij and aij take on the values 0, 1, or 2.) Assuming an adiabatic system and no intraparticle temperature gradients, the differential heat balance equations for gas and solid phases are written as
and (5)
20,
ws
- Pg)TDtaU,
(10)
Neglecting the entrance effect yields the maximum initial velocity, corresponding to the pseudo-steady-state assumption; i.e., the relative velocity of gas and solid is assumed equal to the terminal velocity in a multiparticle system, as given by eq 9 in each reactor segment. The VPTR model is thus divided into two submodels, one neglecting the entrance effect and the other accounting for it as described above. The fluid mechanical parameters of eqs 7-9 are determined as follows. The drag coefficient for a single sphere particle in a single-particle system can be expressed as (Yang, 1973) 24 CDs = -[1 + 0.150(Re,)0.687] for Re, < 1000 (11) Re, where (12) Re, = P.bg - u,)d,/rcg The solid friction factor can be expressed as (Yang, 1978) for ug/ut > 1.5
f, = 0.0126€3
where A = 6(1 - t)/d, (6) is the surface area of particles per unit bed volume. The pressure gradient is calculated by solving the following differential momentum balance equation for the gasaolid mixture, accounting for the contributions due to
(13) f, = 0.0410-
where
for ug/ut
< 1.5
Ind. Eng. Chem. Res., Vol. 31, No. 10,1992 2267 Ret = d,utP,/cc,
(15)
ut is the terminal velocity in a single-particle system and is given by (Yang, 1973) 0.153d,1~14g0~71(p, - p,)O.'l Ut = for Re, < 1000 (16) d.43p,0.29
The gas friction factor is found as follows (Fujita, 1976): f, = 0.0791(Re,)4"5 for Re, Ilo5 (17) l/f,'/2 = 4.0 log (fg1/2Re,)- 0.4 for Reg > lo5
(18)
where Re, = PgUgDt/Pg
(19)
Finally, the VPTR model is completed by incorporation of a suitable correlation for gas-particle heat-transfer coefficients expressed in terms of the Nusselt number, Nu. Assuming that the gas and solid phases are in plug flow in the region of dilute phase transport, the gas-particle model developed by Nelson and Galloway (1975) was selected for thia study. Their model was based on low-temperature studies without chemical reactions and neglecting contributions from thermal radiation. It was selected as the most appropriate model for this work in the absence of other data. Thus,
Nu =
\-
104 10-5 10-4 10-3 10-2 10-1
100 1 0 1
102
(20)
c
S ' - tanh 1- (1- 4 1 / 3
Table 11. System Parameters Employed for Numerical Simulation of the Transport Reactor for Ethane Cracking (See Figures 1-5) fixed param Tgi ("C) 650 Tsi ("C) 900 pti (atm abs) 2.5 Stm/HC (kg of steam/kg of ethane) 0.4 pS (kg/m3) 2375 14.0 (for Figure 1) L (m) 60 (for Figures 2-5) X(C,Hd (%) W H C (kg/s) 9500 varied param d, (rm) 25-400 Sd/HC (kg of solid/kg of ethane) 10-50 D,(m) 0.5-1.0 Uoi (m/s) 5.6-22.5
6
with
Here Fs is the Frossling number, whose value is taken at 0.6. Nelson and Galloway assumed potential flow in the gas phase around stationary fine particles. Therefore, in order to apply the model to the dilute system of suspended fine particles, the superficial particle Reynolds number is given in terms of the relative velocity as (22) ResO= [P,(u, - u,)d,/~,I~ The foregoing equations constitute the VPTR model used in this work. Since the kinetic parameters (time constants) of the radical and molecular species in this system are of different orders of magnitude, the resulting first-order ordinary differential equations are mathematically stiff. Therefore, Gear's method (1971a,b) was selected for numerical integration of the equations. Numerical solution was carried out on a Zenith 386 microcomputer using a program written in FORTRAN. Computing time was between 10 and 20 min per case. Further details of this model are presented elsewhere (Koyama, 1990).
Results and Discussion The VPTR model was first used to evaluate the potential performance of thistype of reactor for thermal cracking of ethane. Realistic values were taken for system parameters, and these are presented in Table 11. Initial simulations were aimed at determining the importance of the entrance effect mentioned earlier. Figure 1compares typical simulated results of the model taking into account or neglecting the entrance effect with the same overall reactor length. In both cases, once the gas
0.701 ' ' ' ' ' 104 10-5 10-4 10-3 10-2 10-1
'
100
'
101
102
Dfstance [ml Figure 1. Typical simulated results of the transport reactor model for ethane cracking: d, = 50 pm; Sd/HC = 20 kg of solid/kg of ethane; D, = 0.6 m; L = 14.0 m.
is combined with the hot inert solids, the gas temperature increases rapidly due to heat transfer through the large interfacial area of the fine particles and reaches a maximum value in a short distance. Gas and solid temperatures decrease gradually thereafter because of the endothermic cracking reactions in the adiabatic reactor. The gas and solid temperature profiles in Figure l a show that the entrance effect model provides for higher heat transfer between gas and solids in the entrance region. The entrance effect model also yields slightly higher ethane conversion at any given distance than the model which neglects the
2268 Ind. Eng. Chem. Res., Vol. 31, No. 10,1992 0
200
100
300
500
400
-
050
040-
020
-
0 IO
-
0.30
~~
6501
10-4
'
""""
10-3
'
""""
10-2
'
".""'
Resldence tlme
lo-'
'"',-I
I00
[SI
Figure 2. Effects of particle diameter on gas temperature along the transport reactor for ethane cracking; pt, = 2.5 atm abs; Sd/HC = 20 kg of solid/kg of ethane; D,= 0.6 m; X = 60%; entrance effect neglected.
entr&ce effect; however, the conversions in both cases are essentially identical at the reactor end (Figure lb). As shown in Figure lc,d, when the entrance effect is neglected, the void fraction and the slip velocity (the relative velocity between gas and solids) are nearly constant throughout the reactor. However, when entrance effects are considered, the solid holdup and the slip velocity are initially very high and decrease toward the steady-state values. This high slip velocity at the entrance region leads to a high particle Reynolds number (from eq 22) and, consequently, a high gas-particle heat-transfer coefficient (from eq 20 and 21). The low void fraction at the entrance region also results in high heat-transfer area per unit volume. These factors combine to yield a higher rate of gas temperature increase when the entrance effect is taken into account. It should be noted, however, that the logarithmic length scale in Figure 1exaggerates the importance of this effect; in fact entrance effects are essentially negligible after 1m of reactor length. Although curves for both cases are presented in some subsequent figures, the effect may generally be considered unimportant for the current range of system parameters. Parametric Sensitivity of the Transport Reactor Model. The effects on ethylene selectivity of particle diameter, solid to hydrocarbon ratio, superficial gas velocity, and reactor diameter were first studied at fixed conversion. Ranges of these parameters considered are indicated in Table 11. Figure 2 shows how gas temperature varies with gas residence time for several particle diameters, neglecting the entrance effect. When the particle diameter is about 50 pm, the gas temperature reaches its maximum value fastest; this maximum is also the highest. An increase in particle diameter increases the slip velocity significantly (eq 8 or 9) and decreases the void fraction slightly (eq 10). Therefore, the particle diameter and slip velocity have the major effect on the gas-particle heat-transfer coefficient (eq 20). Since an increase in the particle diameter increases the Nusselt number (eq 20) but decreases the gas-particle heabtransfer area (eq 6),there is a particle diameter which yields the maximum gas-particle heat-transfer rate per unit bed volume. This turns out to be approximately 50 pm. Figure 3 shows the effect of the particle diameter on the required residence time to achieve 60% conversion and the corresponding ethylene selectivity. The data show that a 50-pm-diameter particle requires the shortsst residence time to achieve 60% conversion and yields the maximum ethylene product. This is because the average gas temperature throughout the reactor is highest when the par-
Neplectlng the entrancs effect 0
75.01
0
"
100
cOnsl&rlnp theentrameffect
"
200
"
300
"
400
1
" 500
Partlcle diameter h m l
Figure 3. Residence time and ethylene selectivity vs particle diameter at 60% conversion in the transport reactor for ethane cracking: pti = 2.5 atm abs; Sd/HC = 20 kg of solid/kg of ethane; D, = 0.6 m.
-
850
u
L
f E
a
800
750
---t
WHC=ZO WHC.30 WHC-40
-C
WHC-50
--O-
Figure 4. Effects of solid to hydrocarbon ratio on gas temperature along the transport reactor for ethane cracking; pti= 2.5 atm abs; d, = 50 Fm; D, = 0.6 m; X = 60%; entrance effect neglected.
ticle diameter is 50 pm as shown in Figure 2. Van Damme et al. (1984)have previously explained that higher reaction temperature or shorter residence time improves the ethylene selectivity in thermal cracking of hydrocarbons. Figure 4 illustrates how gas temperature varies with gas residence time as a function of the solid to hydrocarbon ratio. As the solid to hydrocarbon ratio increases, gas temperature increases rapidly at a l l points within the reactor. The latter result follows because the amount of heat supplied to the reactor increases as the solid to hydrocarbon ratio increases; the increasing rate of heat transfer to the gas follows from the interplay of solid to hydrocarbon ratio, void fraction, and slip velocity with void fraction having the major effect. Figure 5 shows that increasing the solid to hydrocarbon ratio leads to a decrease in the required residence time as well as an increase in ethylene selectivity at 60% conversion. The effects of the gas superficialvelocity and the reactor diameter on the ethylene yield and the required residence time were found to be very small; neither parameter has a significant effect on the slip velocity or the void fraction
Ind. Eng. Chem. Res., Vol. 31, No. 10,1992 2269 1
-
0
2
\
- 1 C
3
0
4
0
5
0
6
a
0.60
u
E
0
0.40
0
Table 111. Comparison of Predicted Yields for the Heater Coil Reactor and the Transport Reactor for Ethane Cracking VPTR heater coil ~~
I
simulation param L (m) D,(m) d, (rm) Sd/HC (kg/kg) T d ("C) Tsi ("C) Tgo("C) P,., (atm abs) dP (atrn) Stm/HC (kg/kg) wHC (kg/h) t
95.00 0.108 680.0
900.0
(8)
X(C,H!) product yield (wt % ) H2
CH4 CZH4 C3Hs 1,3-C,HB c5+
1
0
2
0
3
0
Y
)
5
0
6
0.0 I
0
or consequently on the gas-particle heat-transfer rate. Comparison of the Transport Reactor with a Pyrolysis Furnace for Thermal Cracking of Ethane and Propane. The balance of this paper compares the performance of the VPTR with a typical pyrolysis furnace for cracking ethane and propane. A heater coil model for the pyrolysis furnace was formulated on the basis of that presented by Froment and Bischoff (1979). For ethane cracking, the heat flux profile along the coil was taken to be the same as used by Froment and Bischoff (1979)for a system with four parallel coils. That profile produced reasonable coil wall temperatures (maximum of 906 "C) while resulting in significant conversion of ethane. As will be shown in subsequent tables and figures, the exit temperature reached in the VPTR base case simulation (solid to hydrocarbon ratio of 40) is reasonably close to that calculated for the heater coil. Since conversions were essentially the same, this implies that the net amount of heat added to the reacting gas stream was the same in both cases. It should be noted that the temperature profiles through the reactors are considerably different, owing to the different mechanisms by which heat is transferred to the gas in each case. For propane cracking, the heat flux profile was adjusted so that the conversion achieved was 88.670,a typical value for actual practice. In order to compare product yields (especiaUy ethylene yield), the gas inlet temperature, outlet pressure, steam to hydrocarbon ratio, and conversion in the two reactors were fixed at the same values. It was also assumed that the capacity of the transport reactor was equivalent to that of the four-coil pyrolysis furnace; consequently, the mass flow rate of hydrocarbon for the transport reactor was taken to be 4 times as high as that for each of the individual heater coils in the furnace. The parameters used and results for ethane and propane cracking are shown in Tables I11 and IV, respectively. In the case of the transport reactor model, the particle diameter was fixed at 60 pm but the solid to hydrocarbon
~
4.94 0.60 50
40 680.0 900.0
649.2 1.356 1.634 0.4 2265 0.592 60.81
785.3 1.361 0.039 0.4 9060 0.565 60.79
9060 0.194 60.81
9060
3.793 3.209 47.926 0.653 1.563 3.483
3.645 2.640 48.421 0.691 1.782 3.075
3.865 2.585 49.551 0.516 1.704 2.455
3.870 2.479 50.033 0.441 1.663 2.170
0.2
0.1
~
8.10 0.60 50 30 680.0 900.0 818.8 1.347 0.053
'
0.3 "
0.4 ' '
"
836.8 1.329 0.071 0.4
0.4
0.5 '
0.116 60.79
0.6 '
(
eso
solld to hydrocarbon ratlo Ikg/kgl Figure 5. Residence time and ethylene selectivity vs solid to hydrocarbon ratio at 60% conversion in the transport reactor for ethane cracking: pti = 2.5 atm aba; d, = 50 wm; D,= 0.6 m.
14.00 0.60 50 20 680.0
~~
f
800
n
750
g
-
E
C
700
c
a
e
C
-
1.0
'
J
: . :
Heatercoil reactor VPTR (65.50, MIHC=PO) VPTR (65.50, sd/HC=30)
'
e
:
: . :
'
:
~
60
c
Ez 0
u
40 C
20
0 0.0
0.1
0.2
0.3
0.4
Resldence tlme
0.5
0.6
[SI
Figure 6. Gas temperature, totalgee preaswe, and conversion along the heater coil reactor and the transport reactor with various solid to hydrocarbon ratios for ethane cracking: Td = 680 "C; pto = 1.35 atm abs; X = 60.8%.
ratio was varied from 20 to 40 kg/kg. The gas temperature, total gas pressure, and conversion variation with gas residence time for the heater coil and the transport reactor for ethane cracking are compared in Figure 6. In the case of the heater coil reactor, the gas temperature and the conversion increase almost linearly with residence time, and the outlet temperature is higher
2270 Ind. Eng. Chem. Res., Vol. 31, No. 10, 1992 83 I
Table IV. Comparison of Predicted Yields for the Heater Coil Reactor and the Transport Reactor for Propane Cracking heater VPTR coil
883.5 1.402 1.598 0.40 2450 0.669 88.45
29.80 0.60 50 17 650.0 900.0 793.8 1.407 0.043 0.40 9800 0.797 88.46
20.15 0.60 50 20 650.0 900.0 807.3 1.405 0.045 0.40 9800 0.532 88.44
9.50 0.60 50 30 650.0 900.0 834.9 1.386 0.064 0.40 9800 0.243 88.46
6.60 0.60 50 40 650.0 900.0 849.7 1.363 0.087 0.40 9800 0.165 88.44
1.170 24.177 35.315 19.232 2.450 4.051
1.204 24.340 35.863 18.430 2.944 4.041
1.213 24.347 36.037 18.329 3.006 3.957
1.227 24.362 36.334 18.220 3.134 3.738
1.233 24.357 36.422 18.201 3.211 3.593
95.00 0.108 650.0
-5 s
L
50
a H&W mll resctw
-
60
-
79
-
0
0.2
0.0
0.6
0.4
0.8
Resldence tlme [SI Figure 8. Ethylene selectivity vs residence time at 60.8% conversion in the heater coil reactor and the transport reactor with various solid to hydrocarbon ratios for ethane cracking: Tgi= 680 "C; pto = 1.35 atm abs. 0.2
0.0 42
I .o
0.8
0.4
0
40-
m
5 W
30
-
20
-
-
W 0
c
n
5 Y
81
Trmsport reactor (&=50um)
l Y
0
3
-
m
c c
62
w
I
3gl------
22.0
"L ---t
0
----t
VPTR(Sd/HC=20,45=50) VPTR(Sd/HC=30,45=50) VPTR(Sd/HC=40,ds=50)
c P)
0
20
40
60
0
c
e
o Heater mil reactor
n
Figure 7. Ethylene yield vs conversion in the heater coil reactor and the transport reactor with various solid to hydrocarbon ratios for ethane cracking: Tgi = 680 "C; pt, = 1.35 atm abs; X = 60.8%.
than that of the transport reactor even when the solid to hydrocarbon ratio is 40 kg/kg. The heater coil reactor requires a longer residence time than the transport reactor to achieve 60.8% conversion even at a solid to hydrocarbon ratio of 20 kg/kg. Furthermore, pressure drop in the heater coil reactor is significant because of the very high gas linear velocity and the high coil length necessary for adequate gas-wall heat transfer. The ethylene yield versus the ethane conversion in both reactors is compared in Figure 7. It is clear that the transport reactor yields more ethylene product than the heater coil reactor at any given conversion. As shown in Figure 8, although the transport reactor can achieve a shorter residence time to improve the ethylene selectivity, even at the same residence time the ethylene selectivity of the transport reactor is 0.75% higher than that of the heater coil reactor. This occurs because of the higher average gas temperature and much lower gas pressure drop in the transport reactor. Figure 9 shows ethylene and propylene selectivity versus the residence time at 88.5% conversion for propane cracking. Although the ethylene selectivity of the transport reactor for propane cracking is less sensitive to the residence time than in the case of ethane cracking, the ethylene selectivity of the heater coil reactor is still 0.70%
20.0-
0
Converslon [XI
VPTR (bs-50um) 19.01
'
0.0
a
0.2
'
I
0.4
'
I
0.6
'
'
'
0.8
Residence time Is1 Figure 9. Ethylene and propylene selectivitiesvs residence time at 88.5% conversion in the heater coil reactor and the transport reactor with various solid to hydrocarbon ratios for propane cracking: Tgi = 650 "C;pto = 1.4 atm abs.
less than that of the transport at the same residence time. On the other and, propylene selectivity of the transport reactor is somewhat less sensitive to the residence time than ethylene selectivity. Furthermore, propylene selectivity of the heater coil reactor is 0.95% higher than that of transport reactor at the same residence time. This difference in performance is due to the chemistry of the propane cracking reactions. The reactions in this scheme which related to propylene production in propane cracking are as follows (Sundaram and Froment, 1978): reaction 1% Af/E (a) C3H8 + H' l-C3H7*+ Hz 11.0/9.7 (b)
C3Hs
+ H'
--
2-C&*
+ Hz
11.0/8.3
Ind. Eng. Chem. Res., Vol. 31, No. 10, 1992 2271 Propane produces two different propyl radicals through the hydrogen abstraction reaction (reactions a and b). Though 2-C3H7*contributes only to propylene formation (reaction e), 1-C3H,*can produce both ethylene and propylene (reactions c and d, respectively). Because of the magnitudes of frequency factors and activation energies, ethylene production via reactions a and c is favored compared to propylene production via reactions a and d as well as reactions b and e. This behavior is somewhat more pronounced at higher temperatures. As a result, in the case of propane cracking, the transport reactor will produce more ethylene but less propylene than the pyrolysis furnace. These calculated results have demonstrated that the VPTR offers a much more efficient means for providing the heat necessary to initiate thermal cracking of ethane and propane gases than the more conventional pyrolysis furnace coil reactor. It is possible with the VPTR to achieve comparable conversions at much lower residence time or, conversely, to increase the olefin production rate in the reactor. Clearly, it is the temperature profile which is attained in the reacting gas streams that determines the conversion-selectivity patterns found, and this profile is directly dependent on the means by which heat is added. The VPTR makes it possible to raise the gas temperature very rapidly, as opposed to the essentially linear increase with residence time which is attainable in the heater coil reactor, and thus to take advantage of the small differences in kinetic parameters which favor olefin production over other reaction paths. While the net differences in selectivities may not be great, even improvements on the order of 1or 2 % may be significantin the large-scale production characteristic of thermal cracking operations. Another advantage for VPTR operation suggested earlier is the freedom from the problems associated with coke deposition on reactor walls. While the present study did not consider that effect explicitly (coke formation reactions were not considered important in the ethane and propane cracking reactions studied here), coke formation will be of significance in thermal cracking of large hydrocarbon molecules. Consideration of corresponding reactions as well as the simulation of an entire reactor-regenerator system poses a more complex and difficult problem than that studied in this work. It would, however, be a natural extension of the present investigation. Finally, the possibility of incorporating catalytically active solids which will not only serve as heat carriers but which will also promote desired reactions enhances the potential merit of the VPTR approach to thermal and catalytic processing of other reactive gases. Further applications of the VPTR exploiting these advantages could be of considerable practical interest.
Conclusions These model studies demonstrated the following: 1. The vertical pneumatic transport reactor can achieve high conversions at shorter residence times and with improved ethylene selectivity in ethane and propane cracking than the pyrolysis furnace. 2. The solid particle diameter and the solid to hydrocarbon ratio have the greatest effect on the performance of the transport reactor. The optimum particle diameter is approximately 50 pm. 3. For small-diameter particles and small solid to hydrocarbon ratios, particle accelerationeffeds at the reactor entrance may be neglected without introducing significant errors.
Nomenclature A = gas/solid interfacial area per unit reador volume (m2/m3) Af = frequency factor (s-l or m3/(kg-mol.s)) ai,= reaction order of component j in the ith reaction CDs= drag coefficient on a single particle in a single-particle
system C , = molar concentration of component j (kg-mol/m3) C,, = heat capacity of gas mixture and solid (kcal/kg.K) D, = tube diameter (m) d, = solid particle diameter (m) E = activation energy (kcal/mol) Fi = gas molar flow rate of species i (kg-mol/s) Fs = Frossling number f,, f, = friction factor of gas and solid g = gravitational acceleration (m/s2) AHi = heat of ith reaction (kcal/kg-mol) h,, = gas-solid heat-transfer coefficient (kcal/(m*.s.K)) ki = rate constant of ith reaction (s-l or m3/(kg-mol.s)) L = reactor length (m) Nu = Nusselt number expressed as h,,d,/A, Pr = Prandtl number expressed as Cpgpg/Ag pt = total gas pressure (atm abs) R = gas constant (m3.atm/(kg-mol.K)) Re, = gas Reynolds number expressed as pgu.Jlt/pe Re, = particle Reynolds number defined as p,(u, - u s ) d / p gfor a suspended particle system Re, = superficial Reynolds number defined as p,(u, - us)td,/y for a suspended particle system Re, = terminal Reynolds number defined as psu,d,/pg ri = reaction rate for ith reaction (kg-mol/(m3.s)) Sd/HC = solid to hydrocarbon ratio (kg of solid/kg of hydrocarbon) Stm/HC = steam to hydrocarbon ratio (kg of steam/kg of hydrocarbon) T,,T,= temperature of gas and solid (K) t = gas residence time ( 8 ) U,= superficial gas velocity (m/s) u,, u, = linear velocity of gas and solid (m/s) ut = terminal velocity in a single particle system (m/s) WHc = hydrocarbon mass flow rate (kg/h) W,, W , = mass flow rate of gas and solid (kg/s) X = conversion (%) z = axial vertical distance (m)
d,,
Greek Symbols ai, = stoichiometric coefficient of species j in the ith reaction t = void fraction A, = heat conductivity of gas mixture (kcal/(m.s.K)) p, = viscosity of gas mixture (kg/(m*s)) p,, pa = density of gas mixture and solid (kg/m3) Subscripts i = at reactor inlet o = at reactor outlet Registry No. C2&, 74-840;C3H8,7498-6;CzH4,7485-1;C3&, 115-07-1.
Literature Cited Bandrowski, J.; Kaczmarzyk, G. Gas-to Particle Heat Transfer in Vertical Pneumatic Conveying of Granular Materials. Chem. Eng. Sci. 1978, 33, 1303.
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Fan, L.-S.; Satija, S.; Kim, B. C.; Nack, H. Noncatalytic Gas-Solid Reactions in a Vertical Pneumatic Transport Reactor. AZChE J. 1984, 30, 21. Froment, G. F.; Bischoff, K. B. Chemical Reactor Analysis and Design; Wiley: New York, 1979; p 410. Fujita, S. Kagaku-Kougaku Z, 2nd ed.; Iwanami-Zensho: Tokyo, 1976; p 44. Cartside, R. J.; Ellis, A. F. Thermal Regenerative Cracker: A Development Update. Chem. Eng. h o g . 1983, 79,82. Gear, C. W. The Automatic Integration of Ordinary Differential Equations. Commun. ACM 1971a, 14, 176. Gear, C. W. Numerical Initial Value Problems in Ordinary Differential Equations; Prentice-Hall: Englewood Cliffs, NJ, 1971b. Goosens, A. G.; Dente, M. Simulation Program Predicts Olefin Furnace Performances. Oil Gas J. 1978, Sept 4,89. Goosens, A. G.; Dente, M.; Ranzi, E. Improve Steam Cracker Operation. Hydrocarbon Process 1978, Sept, 227. Jepson, S. C. Computer Simulation of a Non-Isothermal Vertical Pneumatic Transport Reactor for Naphtha Pyrolysis. M.S. Thesis, Northwestern University, Evanston, IL, 1986. Kato, K.; Onozawa, I.; Noguchi, Y. Gas-Particle Heat Transfer in a Dispersed Bed. J. Chem. Eng. Jpn. 1983, 16, 178. Koyama, H. Modeling of Thermal Cracking of Ethane and Propane in A Non-Isothermal Vertical Pneumatic Transport Reactor. M.S. Thesis, Northwestern University, Evanston, IL, 1990. Nelson, P. A,; Galloway, T. R. Particle-to-Fluid Heat and Mass Transfer in Dense Systems of Fine Particles. Chem. Eng. Sci. 1975, 30, 1. Paraskos, J. A.; Shah, Y. T.; Mckinney, J. D.; Carr, N. L. A Kine-
matic Model for Catalytic Cracking in A Transfer Line Reactor. Znd. Eng. Process Des. Dev. 1976, 15, 165. Pratt, K. C. Catalytic Reactions in Transport Reactors. Chem. Eng. Sci. 1974, 29, 747. Ross, L. L. Pyrolysis furnace design: Conventional and Novel. In Pyrolysis: Theory and Industrial Practice; Albright, L. F., Crynes, B. L., Corcoran, W. H., Eds.; Academic Presa: New York, 1983. Shaikh, A. A.; Carberry, J. J. Model of Isothermal Transport-line (Riser) and Moving-bed Catalytic Reactor. Chem. Eng. Res. Des. 1984, 62, 387. Sundaram, K. M.; Froment, G. F. Modeling of Thermal Cracking Kinetics. 3. Radical Mechanisms for the Pyrolysis of Simple Paraffii, Olefins, and Their Mixtures. Znd. Eng. Chem. Fundam. 1978, 17, 174. Van Damme, P. S.; Willems, P. A,; Froment, G. F. Temperature, not time, controls steam cracking yields. Oil Cas J. 1984, Sept 3,68. Varghese, P.; Varma, A. Catalytic Reactions in Transport-line Reactors. Chem. Eng. Sci. 1979, 34,337. Yang, W.-C. Estimating the Solid Particle Velocity in Vertical Pneumatic Conveying Lines. Ind. Eng. Chem. Fundam. 1973,12, 349. Yang, W.-C. A Unified Theory on Dilute Phase Pneumatic Transport. J. Powder Bulk Solids Technol. 1977, I, 89. Yang, W.-C. A Correlation for Solid Friction in Vertical Pneumatic Conveying Lines. AZChE J. 1978,24, 548. Received for review June 12, 1992 Accepted July 2, 1992
Partial Oxidation of Methane: The Role of Surface Reactions Daniel J. Thomas, Ren6 Willi, and Alfons Baiker* Department of Chemical Engineering and Industrial Chemistry, Swiss Federal Institute of Technology, ETH-Zentrum, CH-8092 Zurich, Switzerland
The gas-phase partial oxidation of methane with molecular oxygen, to methanol and formaldehyde, has been studied in a flow reactor at 2 MPa and 400-650 O C . The influence of reactor surface area on the reaction was examined both experimentally and by using the model developed by Bedeneev et al. Increasing the surface to volume ratio in a quartz-lined reactor was found to significantly decrease both the reaction rate and the selectivity to methanol. Under favorable conditions, combined methanol/formaldehyde selectivities of over 60% were obtained, although at methane conversions of below 1%. The implications of these results for both previous and future work are discussed.
Introduction The partial oxidation of methane to methanol and/or formaldehyde has considerable potential for the utilization of vast natural gas fields in remote areas of the world. The method offers not only ease of transportation (due to the reduction in volume) but also increases the range of subsequent applications for further processing. The number of recent publications in this area reflects renewed interest in this reaction as an alternative to the two-stage steamreforming route to methanol. The partial oxidation reaction is, potentially, a simpler and more energy-efficient process than the steam-reforming route. Only a single exothermic step is required, as opposed to the highly endothermic steam-reforming reaction followed by exothermic methanol synthesis. Edwards and Foster (1986) reported that methane partial oxidation should be economically favorable over steam reforming, if methanol yields in excess of 77% at pass conversions above 4 % can be achieved. Although some workers have reported methanol selectivities in this region (Hunter et al., 1984,1990),these results have proven difficult to reproduce in other laboratories. It is well established that homogeneous gas-phase reactions are very
sensitive to temperature distribution and flow configuration, and it has been suggested by Brown and Parkyns (1991) that this may be the cause of the large discrepancies in reported methanol yields. Considerable attempts have been made to find suitable catalysts for methane partial oxidation. These have generally led to relatively higher formaldehyde selectivities which drop off rapidly with increasing conversion (e.g., Ahmed and Moffat (1988) or Kastanas et al. (1988)). There is a growing awareness that noncatalytic yields are generally comparable to or better than those for the catalytic reaction. This has led to several recent publications examining the influence of homogeneous gas-phase reactions on catalytic systems. While catalytic systems are clearly affected by homogeneous gas-phase reactions, the reverse is certainly also true: the homogeneous system is affected by surface reactions. This applies even in the case of "inert" surfaces such as quartz. Baldwin et al. (1991) found that the presence of inert packing effectively quenched the radical reactions, resulting in lower product yields. However, a quantitative investigation of the effect of reactor surface/volume ratio on the reaction has yet to be carried out.
Q888-58S5/92/2631-2272$03.oo/o 0 1992 American Chemical Society
