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Ind. Eng. Chem. Res. 2003, 42, 6-13
Partial Hydrogenation in an Upflow Fixed-Bed Reactor: A Multistage Operation for Experimental Optimization of Selectivity Frank Stu 1 ber*,† and Henri Delmas‡ Departament d’Enginyeria Quı´mica, ETSEQ, Universitat Rovira i Virgili, Paisos Catalans 26, 43007 Tarragona, Catalunya, Spain, and Laboratoire de Ge´ nie Chimique UMR, CNRS 5503 ENSIGC, 18 Chemin de la Loge, 31078 Toulouse Cedex 4, France
The selectivity performance of an upflow fixed-bed reactor was optimized for the exothermic consecutive hydrogenation of 1,5,9-cyclododecatriene (CDT) to cyclododecene (CDE) over 0.5% Pd/alumina. The influence of the main operating variables (temperature, hydrogen pressure, and liquid and hydrogen flow rates) on global hydrogen conversion and CDE selectivity was studied and analyzed. The reactor performance obtained for a pellet catalyst was shown to be sensitive to external mass transfer of the gaseous reactant and backmixing in the liquid phase. The surface concentration of hydrogen was found to be the key parameter because of its opposite effect on the rate of hydrogenation and selectivity. Increasing the gas velocity provided a higher hydrogen concentration at the catalyst surface, enhancing the global rate of hydrogenation, although an increase in the global hydrogen conversion was always related to a reduction of CDE selectivity. When the liquid velocity was decreased, backmixing became important in the liquid phase, also resulting in a loss of selectivity. To improve the reactor performance in terms of selectivity, process staging and hydrogen dilution by an inert gas were tested. Both, but in particular the split-up of the process into several stages, proved to be successful in obtaining high yields of CDE up to 90%, required for industrial application. Introduction Multiphase catalytic reactors involving gas, liquid, and solid phases are widely encountered in industrial practice. The relevant fields of application cover petrochemical (hydroprocessing), fine chemical and pharmaceutical (hydrogenation), biochemical, and more recently environmental (oxidation) processes.1,2 The most common reactor type used in these industrial applications is undoubtedly the trickle-bed reactor with downflow of gas and liquid phases, although different modes of operation, such as gas-liquid upflow, are also becoming important.1,3 The principal question may arise of whether to use a downflow or an upflow mode of operation to obtain high conversion and selectivity in gas-liquid reactions catalyzed by a solid. The former mode is characterized by a continuous gas phase, and lower liquid holdup may enhance reaction rates by direct gassolid contact (partial wetting of the catalyst). The latter mode may have advantages with respect to uniform liquid distribution (higher liquid holdup), temperature control, and removal of deactivating byproducts, resulting in a better selectivity performance for consecutive reactions. Previous research work reported in the literature has provided the answer only for a specific reaction system, mostly restricted to single reactions under isothermal conditions.4-11 Khadilkar et al.11 examined the previously reported studies and concluded that, under conditions of liquid-phase reactant limitation, the upflow reactor outperforms the trickle-bed reactor because of its more efficient liquid-solid contact. On the other hand, detailed studies of reactor performances under * Corresponding author. Tel: 00 34 977 559671. Fax: 00 34 977 559667. E-mail:
[email protected]. † Universitat Rovira i Virgili. ‡ CNRS 5503 ENSIGC.
nonisothermal conditions wherein complex multistep reaction has been considered, are less frequent for trickle-bed reactors12-15 and rare for upflow fixed-bed reactors.14-18 More research work is necessary in this field because an important issue in the design and scaleup of a fixed-bed reactor is the control of temperature and selectivity in exothermic multistep reactions. In particular, rigorous studies with respect to the selectivity performance in the upflow mode are scarce for exothermic multistep hydrogenation reactions.19-21 A comparison of the performances of up- and downflow fixed-bed reactors for the consecutive hydrogenation of pyrolyzed gasoline,22 phenylacetylene,19 and butadiene feedstocks,20 suggested that the upflow bed reactor is advantageous, matching higher selectivity, longer catalyst life, and lower residue production. Recently, the selective hydrogenation of 1,5,9-cyclododecatriene (CDT) to cyclododecene (CDE) was studied in up- and downflow operation modes over a shell-type 0.5% Pd/alumina catalyst.14,15 The most important conclusion of these works is that a direct comparison of the reactor performance was impossible with a pure CDT feed. Tricklebed operation using a pure CDT feed caused hot-spot formation with an almost instantaneous temperature rise up to 573 K due to direct gas-solid hydrogen mass transfer and very poor heat transfer in the trickle bed in the low interaction regime. To enable a comparative study, the liquid reactant feed and also the catalyst was diluted. Additionally, the reactor wall temperature was decreased to values ranging from 373 to 400 K. Even in these mild operating conditions, a slightly higher rate of hydrogenation combined with a significantly superior CDE selectivity was observed in the experiments with upflow operation mode. Thus, the aim of this work is to present a systematic experimental study of the reactor performance of an upflow fixed-bed reactor for the hydrogenation of CDT
10.1021/ie020466l CCC: $25.00 © 2003 American Chemical Society Published on Web 11/16/2002
Ind. Eng. Chem. Res., Vol. 42, No. 1, 2003 7
using a pure CDT reactant feed and an undiluted catalyst bed of 0.5% Pd/alumina. Experiments have been carried out over a wide range of operating conditions, including process staging and hydrogen dilution, to optimize the CDE selectivity to high molar concentrations (>90%) required for industrial application. Experimental Section Kinetics of CDT Hydrogenation. The kinetics of this reaction has been investigated before, and more details can be found in the related papers.18,23 Catalytic hydrogenation experiments were carried out in a stirred autoclave at different pressures (0.15-1.2 MPa) and temperatures (413-453 K) using two catalyst sizes to determine the intrinsic reaction kinetics: cylindrical pellets of alumina (3.1 mm in diameter) coated with palladium metal to a depth of 250 µm (Degussa, E263/ D, 0.5% Pd) and pellets crushed to a mean diameter of less than 10 µm. In a typical experiment, a known amount of catalyst (10 g in the case of catalyst pellets) along with 250 g of CDT was charged in the reactor and flushed with nitrogen at room temperature. The heating was then started and, once the reaction temperature was reached, the system was pressurized with hydrogen to the desired pressure, and finally the reaction was initiated by turning on the stirrer. During the course of the reaction, liquid samples were regularly taken and analyzed by a HP 5890 gas chromatograph equipped with a HP-FFAP capillary column. As many as 20 peaks appeared on the gas chromatogram. By mass spectroscopy, 15 peaks have been identified as isomers of the four hydrocarbons. Five peaks can be attributed to the CDT, seven peaks to the CDD, two peaks to the CDE, and one peak to the CDA. The area of the five remaining peaks never exceeded 0.2% of the total area, and analysis precision should not be influenced when ignoring these peaks. A simplified reaction scheme was established lumping the respective isomers: H2
H2
H2
1
2
3
CDT 9 8 CDD 9 8 CDE 9 8 CDA r r r
(1)
and rate equations of the Langmuir-Hinshelwood type were found to correctly represent the experimental data:
ri )
mcatkiKjCjCH2Ri 3
1+
(2)
KjCj ∑ j)1
where i is the reaction number (i ) 1-3) and j the respective hydrocarbon (j ) 1, CDT; 2, CDD; 3, CDE). Upflow Fixed-Bed Reactor Setup. The selective hydrogenation was carried out in a pilot-scale packedbed reactor (0.026 m inner diameter and 1.5 m height) with a cocurrent gas-liquid upflow mode and an external cooling device. The reactor tube was filled with 0.71-0.73 kg of cylindrical catalyst pellets as used in the kinetic study (Degussa E263/D, 0.5% Pd). A flexible grid fixed by a tensed spring was placed at the top of the reactor to prevent catalyst particle displacement during operation. The catalyst bed was regularly replaced after a fixed period of use, and prior to new experiments, the catalyst activity and selectivity were tested. Figure 1 shows a schematic diagram of the
Figure 1. Experimental setup of the upflow catalytic fixed-bed reactor: (1) double jacketed reactor, (2) gas-liquid separator, (3) product receiver, (4) liquid preheater, (5) feed liquid pump, (6) gas flowmeters, (7) liquid sample valves, (8) liquid feedstock, (9) heat exchanger, (10) expansion tank, (11) heat exchanger bypass, (12) three-way valve, (13) cooling oil pump, (14) cooling oil heater. Dashed lines: control circuits.
experimental apparatus consisting of the reaction circuit with gas (H2 and N2) and liquid feed supply, the packedbed reactor, a gas-liquid separator with a pressure regulator, and a product storage tank and the heating/ cooling circuit with a double vessel jacket enclosing the reactor tube, a countercurrent heat exchanger with bypass regulated by a three-way valve, a cooling oil pump, and heater. To measure axial temperature and concentration profiles, five thermocouples and five liquid sample valves were attached along the reactor tube. One of the temperature probes contained three thermocouples radially distanced to check the presence of radial temperature gradients. The temperature in the reactor was controlled by means of the cooling oil flowing at a high rate in the double jacket at a nearly constant temperature. The reactor pressure was maintained constant by continuous pressure release at the gas outlet. Both temperatures and pressures (inlet and outlet) were monitored using a microcomputer data acquisition system. Gas and liquid flow rates were adjusted by two calibrated flowmeters and a liquid pump (maximum capacity of 6 × 10-3 m3/h) placed on a balance. Liquid-phase samples were analyzed by gas chromatography (HP5890) using a HP-FFAP capillary column. The other conditions of the GC were an injector temperature of 473 K, a column temperature of 383 K, and a detector (flame ionization, FID) temperature of 523 K. Steady-State Experiments in a Fixed-Bed Reactor. In each experimental run, steady-state conditions were reached following a careful heating procedure of
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Figure 2. Outlet molar fraction of hydrocarbons versus hydrogen velocity (P ) 0.6 MPa, TOil ) 433 K, vL ) 10-3 m/s, and vH2,stoic ) 0.115 m/s).
the reactor. The reactant liquid flow was started after adjusting the flow rate by the liquid pump and preheated to the desired inlet temperature. Before entering the flushed and depressurized reactor, gas and liquid reactants were mixed in an inert fixed bed to homogenize the temperature of phases and achieve gas-liquid equilibrium. The temperature in the reactor was set by maintaining the temperature of the cooling oil at the required value. At temperatures of about 80 K lower than the desired operating temperature, gaseous hydrogen was progressively fed to the mixer unit both to slowly start the reaction and to build up the fixed operating pressure. In this way, the pressure reached stable values within 10 min, while axial and radial temperatures required about 30-40 min. However, regularly withdrawn and analyzed liquid samples indicated stable outlet concentrations only after approximately 60-90 min, depending on the liquid flow rate. Then, the influence of the temperature, pressure, and liquid and gas flow rates on the steady-state reactor performance was investigated by increasing the hydrogen flow rate at a constant pressure, cooling oil temperature, and liquid flow rate. Important concentration differences were observed by the analysis of reactants and products in the exit stream as illustrated in Figure 2 for P ) 0.4 MPa, TOil ) 433 K, and vL ) 10-3 m/s. Hydrogenation experiments with pure CDT carried out in a trickle-bed reactor confirmed thermal instability during operation (hot spots), even in a diluted catalytic bed.14 In the upflow mode, however, it was possible to operate the reactor without any risk of runaway at the selected operation conditions. From monitoring axial temperature profiles during the experiments (see Figure 3), a criterion for thermal stability was to avoid bed temperatures higher than 458 K combined with overstoichiometric hydrogen flow rates that promoted local hot-spot formation. In the experimental runs carried out to optimize selectivity, the effects of hydrogen dilution as well as process staging were studied. Hydrogen was diluted at different flow rates of inert nitrogen measured by a second gas flowmeter. For simulating experimentally a second reactor stage, the required inlet liquid feedstock was obtained by collecting the exit mixture of a first stage. The experiments for a second stage were done using different inlet liquid compositions characterized by high CDE selectivities and small molar fractions of the initial reactant CDT. The whole set of experimental conditions is summarized in Table 1.
Figure 3. Axial profiles of the reactor temperature for three different cooling oil temperatures (P ) 0.4 MPa, vL ) 10-3 m/s, and vH2 ) 0.15 m/s). Table 1. Experimental Conditions of an Upflow Catalytic Fixed-Bed Reactor Study (First and Second Reactor Stage) operating conditions
first stage
second stage
cooling oil temperature (K) reactor pressure (MPa) liquid velocity (m/s) hydrogen velocity (m/s) nitrogen dilution ratio reactor data (Di/Da) reactor (cm/cm) (Di/Da) double envelope (cm/cm) reactor length (m) bed porosity bed density (kg/m3) bed weight (kg) catalyst cylindrical pellet diameter × length (mm × mm) apparent density (kg/m3) pellet porosity mean pore diameter (Å)
423-438 0.2-0.6 (0.25-1.0) × 10-3 0.04-0.2
443-463 0.15-0.25 (0.9-1.2) × 10-3 0.04-0.8 1-3
2.6/2.9 4/5 1.5 0.36 800-830 0.71-0.73 0.5 wt % Pd/Al2O3 3 × 3.2 1300 0.56 90
Experimental Results and Discussion Upflow Fixed-Bed Reactor Performance (First Reactor Stage). The influence of the operating parameters (P, Toil, vL, and vH2) on the reactor performance was studied in the range of operating conditions that provided stable reactor behavior (see Table 1). For a discussion of the results, experimental raw data were converted to global hydrogen conversion, ΩH2 (or global rate of hydrogenation, R), and CDE selectivity, S. ΩH2 is defined as the ratio of the actual hydrogen conversion to the amount of hydrogen required for complete conversion of CDT to the desired CDE (0 < ΩH2 < 1.5):
ΩH2 ) 0.5XCDD + XCDE + 1.5XCDA
(3)
where X is the molar fraction of the respective hydrocarbon CDD, CDE, and CDA. A ΩH2 value of 1 ideally means that all CDT is converted to CDE without forming any CDA. According to the work of Bond and Wells24 on the hydrogenation of alkynes and dienes, the term selectivity denotes the extent to which CDT will yield CDE compared to CDA; thus,
S)
XCDE XCDE + XCDA
(4)
Ind. Eng. Chem. Res., Vol. 42, No. 1, 2003 9
Figure 4. Effect of the hydrogen velocity on global hydrogen conversion at different pressures (TOil ) 433 K, vL ) 10-3 m/s, and vH2,stoic ) 0.115 m/s for complete CDE formation).
Figure 5. Effect of the hydrogen pressure on global hydrogen conversion at different hydrogen velocities (P ) 0.4 MPa, TOil ) 433 K, vL ) 10-3 m/s, and vH2,stoic ) 0.115 m/s for complete CDE formation).
This particular selectivity criterion was selected because the main interest of this work was to obtain a maximum of CDE (XCDE > 90%), preventing at the same time complete hydrogenation to CDA (XCDA < 7% was requested). Effect of the Hydrogen Velocity. Figure 4 illustrates the effect of the hydrogen velocity on the global hydrogen conversion at various pressures, a liquid velocity of 10-3 m/s, and a cooling oil temperature of 433 K. In all runs, ΩH2 was found to increase almost linearly with an increase in the hydrogen velocity. The strong dependence was due to a linear increase in the gas-liquid and liquid-solid mass-transfer coefficients with the gas velocity for the upflow operation mode.25 On the other hand, only the use of the highest hydrogen pressure of 0.6 MPa and overstoichiometric hydrogen velocities (vH2 > 0.115 m/s) resulted in required ΩH2 values close to unity (see Figure 4). For example, at 0.6 MPa and a liquid velocity of 10-3 m/s, the applied hydrogen velocity of 0.15 m/s that gave adequate conversions corresponded to a hydrogen excess of 35% with respect to complete conversion of CDT to CDE. However, overstoichiometric hydrogen flow rates along with higher hydrogen pressure are detrimental for the thermal stability of the reactor, as pointed out before. Effect of the Hydrogen Pressure. The effect of the hydrogen pressure on the global hydrogen conversion is shown in Figure 5 for various gas velocities, a liquid velocity of 10-3 m/s, and a cooling oil temperature of 433 K. As expected, the global hydrogen conversion rose linearly with increasing pressure for all hydrogen velocities. According to Henry’s linear law of absorption
Figure 6. Effect of the cooling oil temperature on global hydrogen conversion at different hydrogen velocities (P ) 0.4 MPa, vL ) 10-3 m/s, and vH2,stoic ) 0.115 m/s for complete CDE formation).
and the kinetic law for the consecutive hydrogenation of CDT, showing first order with respect to hydrogen in the pressure range studied, an increase in the pressure proportionally increased the dissolved hydrogen concentration in the liquid phase and thereby the global rate of hydrogenation. Hence, both the hydrogen velocity and hydrogen pressure must be considered as key parameters for the tuning of the reactor performance. Effect of the Cooling Oil Temperature. The highest reactor temperatures were observed at the highest liquid velocity of 10-3 m/s and overstoichiometric hydrogen velocities (0.15 m/s). From the bed temperature profiles (see Figure 3), the mean reactor temperatures at these conditions were determined to be always about 15 K higher than the applied cooling oil temperature. The effect of the cooling oil temperature on the global hydrogen conversion was found to be weak, as demonstrated in Figure 6 for different gas velocities, a liquid velocity of 10-3 m/s, and a pressure of 0.4 MPa. Although raising the oil temperature about 15 K led to a similar ∆T of the mean temperature inside the catalyst bed, the global hydrogen conversion only increased by about 20%. Apparently, and in agreement with the observed effect of the gas velocity on global hydrogen conversion, the reaction was limited by external mass transfer (of the gaseous reactant), which was found to depend mildly on temperature for the present reaction system.18 Effect of the Liquid Velocity. The global rate of hydrogenation increased slightly with the liquid velocity at three gas velocities, a pressure of 0.4 MPa, and a cooling oil temperature of 433 K (Figure 7). For the operating conditions used in this work, the influence of the liquid velocity on both gas-liquid and liquid-solid mass-transfer coefficients was found to be marginal.25 Also, little effect of the liquid velocity on external masstransfer coefficients was stated in the literature for bubble-flow regime and upflow operation. Thus, the decrease in the global rate of hydrogenation at lower liquid velocity and higher global hydrogen conversion was mainly due to the nonlinear behavior of the kinetics of the consecutive hydrogenation. Selectivity Behavior. In Figure 8, the selectivity of CDE observed in the fixed-bed reactor was plotted against the outlet molar fraction of CDE for all operating conditions studied. The same figure also provided a comparison with respect to the CDE selectivity observed in the stirred batch reactor using the same catalyst pellets placed in a fixed basket.18 In general, the
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Figure 7. Effect of the liquid velocity on the global hydrogenation rate at different hydrogen velocities (P ) 0.4 MPa, TOil ) 433 K, and vH2,stoic ) 0.115 m/s for complete CDE formation).
Figure 8. Selectivity versus molar CDE concentration observed in a stirred batch reactor (mcat ) 10 g, T ) 433 K, and P ) 0.15 and 1.2 MPa) and upflow fixed-bed reactor (first reactor stage): lines, batch selectivity; points, continuous selectivity.
selectivity of CDE decreased with increasing output molar CDE fraction in both reactor types as expected from related research work on selective hydrogenation reactions.26 For a low molar fraction of CDE up to 2025%, values of CDE selectivity remained practically on a constant level close to unity; thus, the reactors operated with optimal selectivity. A further increase in the CDE yield induced a progressive decrease in selectivity, leading, in the case of the upflow fixed-bed reactor, to a CDE maximum of 71-73% (with 15-18% of CDA), almost 20% below the required value for industrial application. Interestingly, the different operating conditions in the continuous reactor had only a marginal effect on the CDE selectivity, which was found to be similar to that of the batch reactor obtained at a pressure of 0.3 MPa. In the case of the batch reactor, the selectivity was clearly improved by decreasing the pressure (see Figure 8). In the batch reactor, at the chemical pellet regime, the surface concentration of hydrogen was equal to the equilibrium concentration at the gas-liquid interface because of vigorous stirring. The changes in the surface concentration with pressure became significant, and improvement of selectivity must be related to decreasing the concentration of hydrogen on the catalyst surface. In the continuous fixed-bed reactor, the presence of strong external mass-transfer resistance had an effect similar to that of the pressure decrease in the batch system, i.e., sensitively lower concentration of hydrogen at the surface, thereby leading to a better selectivity.
On the other hand, the invarying selectivity related to the change in operating conditions suggested that, besides mass-transfer limitations, reactor hydrodynamics, in particular liquid backmixing at the present low liquid velocities, might also contribute to the selectivity performance of the reactor. To this end, residence time distributions were measured by a conventional tracer method at standard liquid and hydrogen velocities of 0.25 × 10-3-10-3 and 0-0.08 m/s, respectively. The same reaction and reactor system was used, replacing the catalyst pellets by inert nonporous glass spheres. Analysis of the obtained concentration-time profiles allowed estimation of the values of the Peclet number. The Peclet number was found to drop from 50 at the highest liquid velocity (10-3 m/s) to about 10 at the lowest liquid velocity (0.25 × 10-3 m/s). Then, the effect of the liquid flow rate on the product composition was studied by comparing axial concentration profiles at high global hydrogenation conversions (ΩH2 ≈ 1). In addition, one liquid sample valve was directly located upstream of the reactor inlet to detect possible concentrations jumps, a typical effect of backmixing in the liquid phase. Figure 9 shows the axial concentration profiles obtained in experiments for three different liquid velocities, a pressure of 0.4 MPa, and a cooling oil temperature of 423 K. These profiles clearly illustrated the presence of liquid backmixing depending on the value of the liquid velocity. For the highest liquid velocity of 10-3 m/s corresponding to a Peclet number of 50, the liquid phase still behaved like that in plug flow. As the liquid velocity decreased, increasing concentration jumps (of CDT and CDD up to 20%) at the reactor inlet as well as more flat profiles (of CDD, CDE, and CDA) were noted, resulting finally in a decrease in the maximum molar CDE concentration for Peclet numbers in the range of 10-25 (see Figure 9). Hence, the selectivity behavior of the upflow fixed-bed reactor was shown to be the result of the interaction of reaction kinetics, mass-transfer phenomena, and reactor hydrodynamics. To correctly understand the underlying aspects that are responsible for the upflow fixed-bed reactor performance (selectivity), it is necessary to develop a fundamental model accounting for all of the above-mentioned phenomena. Improvement of the Reactor Performance (Selectivity). Any improvement of the reactor performance to reach the required molar CDE concentration of 90% will demand an operating strategy to reduce the observed loss of selectivity for increasing conversions. In batch hydrogenation, optimal selectivity was attained by variable pressure (decreasing the pressure sequence) and careful temperature control.26,27 In the given reactor for continuous hydrogenation processes, the low liquid velocity to be selected for high conversion will result in severe backmixing effects and thereby in low selectivity. Thus, improvement of selectivity could be achieved by reactor staging and even better by a proper choice of operating variables for each stage as stated in the literature.27 To simulate the reactor staging with the existing equipment, partial hydrogenation of CDT was performed to reuse the outlet mixture as the feed of a higher stage. Also, hydrogen dilution by inert nitrogen was tested to enhance (and/ or maintain) the external mass transfer in the catalytic fixed bed and thereby the hydrogenation reactions, providing at the same time a favorable hydrogen deficit at the catalyst surface for optimal selectivity.
Ind. Eng. Chem. Res., Vol. 42, No. 1, 2003 11
Figure 9. Effect of the liquid velocity on axial concentration profiles of hydrocarbons at high global hydrogen conversion (P ) 0.4 MPa, TOil ) 433 K, and ΩH2 ≈ 1).
Figure 10. Outlet concentrations of hydrocarbons versus hydrogen velocity for a second reactor stage (P ) 0.25 MPa, TOil ) 458 K, and vL ) 10-3 m/s).
The experimental runs for a second stage were carried out at high cooling oil temperatures (443-463 K), low pressures (0.2-0.25 MPa), hydrogen velocities up to 0.1 m/s, and a liquid velocity of 10-3 m/s to eliminate negative backmixing effects on selectivity. The initial compositions of the liquid feed selected corresponded to an optimal outlet composition of a first stage, i.e., high initial CDE selectivity of about 0.97 along with a small molar fraction of CDT. Process Staging (Second Reactor Stage). Monitoring of axial temperatures during the experiments for a second reactor stage confirmed quasi-isothermal operation of the reactor. Figure 10 exemplarily illustrates the evolution of the hydrocarbons with increasing hydrogen velocity for a second reactor stage at a pressure of 0.25 MPa, a cooling oil temperature of 458 K, and a liquid velocity of 10-3 m/s. In Figure 11, the enire CDE selectivity data, obtained in a second stage, were reported and compared to the results obtained in a first stage. In general, the same trends were observed in the second stage. The CDE selectivity
Figure 11. Upflow fixed-bed reactor selectivity versus molar CDE concentration for all operating conditions of a first and second reactor stage.
progressively decreased with an increase in the molar CDE concentration and was also found to be insensitive to the changes effected in the operating conditions (see Figure 11). The major gain of a second stage was the notable improvement of the maximum CDE concentration to 83%, although related to a high molar fraction of CDA between 10 and 13% due to overstoichiometric hydrogen flow rates (see Figure 10) necessary to enhance the rates of external mass transfer and hydrogenation. Dilution of Hydrogen. The effect of hydrogen dilution on the reactor performance during a higher stage is shown in Table 2 and Figure 12 for different dilution ratios, a pressure of 0.2 MPa, and a liquid velocity of 10-3 m/s. The experimental data highlighted two opposite effects of hydrogen dilution. As expected, an increasing ratio of hydrogen dilution resulted in respective lower global hydrogen conversions (see Table 2). The partial hydrogen pressure in the gas phase was reduced in the presence of nitrogen, and the amount of hydrogen dissolved in the liquid phase proportionally decreased, thus lowering the global hydrogenation rate,
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Table 2. Effect of the Hydrogen Dilution by Inert Nitrogen on the Reactor Performance (P ) 0.2 MPa, TOil ) 443 K, and vL ) 10-3 m/s) vH2 (m/s) 0.033 0.078 0.078 0.078 0.078 a
Fa
PH 2 (MPa)
0.77 1.22 1.5
(inlet) 0.1 0.1 0.062 0.051 0.046
outlet fraction (mol %) CDT CDD CDE CDA 6.6 3.8 1.4 2.4 2.8 3.0
35.3 22.3 8.8 15.4 18.4 19.1
55.8 69.7 81.8 76.9 74.5 73.6
2.1 3.9 8.7 5.2 4.2 4.2
Ω
S
0.77 0.87 0.99 0.92 0.90 0.89
0.965 0.947 0.903 0.937 0.945 0.945
F ) vN2/vH2 (dilution ratio).
Figure 12. Axial profiles of global hydrogen conversion without dilution (vH2 ) 0.035 m/s) and with dilution of hydrogen by inert nitrogen (vH2 ) 0.075 m/s and FN2 ) 1.22FH2) for a second reactor stage (P ) 0.2 MPa, TOil ) 463 K, and vL ) 10-3 m/s).
which was found to be linear with respect to the surface concentration of hydrogen. The benefit of hydrogen dilution was to enable the use of stoichiometric hydrogen flow rates necessary while operating the reactor with a high selectivity comparable to that obtained at under stoichiometric hydrogen flow rates. In other words, in the case of a stoichiometric hydrogen flow (vH2 ) 0.075 m/s) diluted by nitrogen (FN2 ) 1.22FH2), the entire catalytic bed was available for the hydrogenation reactions as shown in Figure 12 by a linear increase of the global hydrogen conversion up to 0.94. Conversely, at a lower hydrogen flow rate (0.035 m/s) without dilution, the global hydrogen conversion increased more steeply at the bottom of the fixed bed but rapidly leveled out to a lower value of 0.89 at the reactor outlet. It can be assumed that the consumption of hydrogen significantly reduced the gas velocity and thereby the G-L and L-S mass-transfer coefficients of hydrogen. As a consequence, the limiting effect of mass transfer on the overall rate of hydrogenation became even stronger. Moreover, the composition of the outlet mixture with respect to the molar CDE fraction was found to be advantageous in the case of dilution. At a same selectivity value of 0.945, an increase of the molar CDE fraction from 0.697 to 0.745 was observed for runs without dilution (vH2 ) 0.035 m/s) and with dilution (vH2 ) 0.075
m/s and FN2 ) 1.22FH2) as shown in Table 2. Hence, the use of hydrogen dilution should offer a tool to reduce the total number of stages necessary to reach the desired selectivity criterion for industrial application. Optimal Reactor Operation. On the basis of the previous experimental results, a mode of reactor operation was established and tested to perform the selective hydrogenation reactions with the desired selectivity. It basically consisted of a combination of the aspects found to improve the reactor performance: (1) staging of the hydrogenation with a constant liquid velocity of 10-3 m/s, selected as a minimum to avoid backmixing in the liquid phase and a maximum to avoid runaway, and with stoichiometric hydrogen flow rates; (2) dilution of hydrogen at a higher stage to ensure both good selectivity and heat and mass transfer; (3) increase of the cooling oil temperature at higher stages; (4) stagewise decrease of the operating pressure. Following this procedure, six reactor stages were necessary to obtain a final product composition that satisfied the criterion of high CDE selectivity (>90%) and low CDD + CDT (
