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Ind. Eng. Chem. Res. 2009, 48, 1096–1106
Vapor-Liquid Equilibrium Study in Trickle-Bed Reactors Jinwen Chen,*,† Neil Wang,† Fabian Mederos,‡ and Jorge Ancheyta‡ National Centre for Upgrading Technology (NCUT), One Oil Patch DriVe, DeVon, AB, T9G 1A8, Canada, and Instituto Mexicano del Petroleo (IMP), Eje Central Lazaro Cardenas Norte 152, Col. San Bartolo Atepehuacan, Mexico D.F. 07730
Vapor-liquid equilibrium (VLE) in trickle-bed hydroprocessing reactors can significantly change the fluid hydrodynamics and the distribution of reacting species in both the vapor and liquid phases and, ultimately, change the reactor performance. VLE is especially important to pilot-plant studies in which ideal operating regimes (plug flow, full catalyst wetting, absence of reactor wall effects, etc.) are desired to generate reliable, reproducible, and representative data for commercial scale-up and kinetics studies. In this article, we report VLE flash experiments that were conducted in a continuous-flow unit with hydrogen and various petroleum middle distillates under typical hydrotreating conditions to study the relative distribution of the oil in the two phases. The experimental data were further used to evaluate the interaction coefficients, required to perform VLE flash calculations, between hydrogen and hydrocarbon boiling-point pseudocomponents. Furthermore, flow hydrodynamics were predicted in a pilot-plant trickle-bed reactor for hydrotreating two different middle distillate feeds to provide a mapping of operating conditions under which the desired operating regimes could be maintained. 1. Introduction Trickle-bed reactors are widely used in heavy oil upgrading, petroleum refining, chemical/petrochemical, and other industries for three-phase reactions. During heavy oil upgrading and petroleum refining, trickle-bed reactors are used to remove sulfur, nitrogen, oxygen, and metals; to saturate olefins and aromatic rings; and to crack heavy hydrocarbons into light hydrocarbons. These processes, referred to as hydrotreating or hydrocracking, are normally operated under high temperatures and pressures (320-420 °C and 30-120 atm) in the presence of a catalyst and hydrogen. Under such conditions, the reaction system (hydrogen and hydrocarbons) is at vapor-liquid equilibrium (VLE) with both the vapor and liquid phases containing hydrogen and hydrocarbons. In a trickle-bed reactor, under conditions of complete catalyst wetting, only the reacting components present in the liquid phase can come into contact with the catalyst surface and undergo reactions. As temperature and hydrogen flow rate increase and pressure decreases, more hydrocarbons vaporize into the vapor phase, which could significantly change the fluid flow rates and hydrodynamics, as well as the concentrations of various reacting species in the two phases. For pilot-plant studies, where a relatively small reactor is used, VLE becomes especially important because any discrepancies from the desired ideal operating regime (plug flow, full catalyst wetting, absence of reactor wall effects, etc.) could result in unreliable and unrepresentative data for commercial scale-up and for kinetics studies. Therefore, it is necessary to study, understand, and correctly predict the VLE in hydroprocessing reactors and, furthermore, to predict the operating regimes under which pilot-plant studies can generate reliable, repeatable, and representative results. Only a few studies have been published considering VLE in three-phase reactors,1-5 with even fewer accounting for VLE effects on petroleum hydroprocessing.6,7 Experimental data on the VLE of hydrogen-petroleum fraction systems obtained * To whom correspondence should be addressed. Tel.: (780) 9878763. Fax: (780) 987-5349. E-mail:
[email protected]. † National Centre for Upgrading Technology (NCUT). ‡ Instituto Mexicano del Petroleo (IMP).
under commercial hydroprocessing conditions are rarely published in the open literature. Instead, systems consisting of hydrogen and a single hydrocarbon or a simulated hydrocarbon mixture are often studied.8-12 Such studies were briefly reviewed in our previous article.13 The virtual lack of research on VLE during petroleum hydroprocessing is probably due to the significant difficulty of experimental measurements (or simulation) and the poor definition of the problem resulting from the complexity of the system. In addition, high temperature and pressure and sample characterization and analysis add more difficulties to such experiments. In this study, by using the calibrated flash calculation program, we predicted VLEs in middle distillate pilot-plant hydrotreaters under various operating conditions. The calculated VLE results were further used to predict the flow dynamics in the hydrotreaters to investigate whether they were operated under the desired operating regime (plug flow, full catalyst wetting, absence of reactor wall effects). Detailed estimations and discussions of these hydrotreater operating regimes are discussed with regard to two typical but different middle distillate feeds under various operating conditions and parameters. 2. Experimental Section 2.1. VLE Experiments. A continuous-flow VLE cell was used to conduct the experiments. The details of the experimental setup, procedures, and validation of experimental data can be found in a previous report.13 Therefore, only a brief description is provided here. The oil feed was pumped from a feed tank and then mixed with hydrogen (99.99%) under the desired pressure. The resulting mixture passed through a stainless steel coil preheater before entering the VLE cell, where it was allowed to flash. Both the coil and the VLE cell were situated in a temperaturecontrolled furnace to ensure a constant flash temperature. The cell (i.d. of 2.54 cm and height of 30 cm) was arranged as a high-pressure, high-temperature phase separator in which a twopoint thermocouple was positioned at the center to measure the temperatures of the liquid and vapor phases separately. In the VLE cell, the liquid level was controlled and maintained
10.1021/ie8006006 CCC: $40.75 2009 American Chemical Society Published on Web 09/04/2008
Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1097 Table 1. Physical Properties and Chemical Compositions of the Materials Used in VLE Experiments mixture of LCO and white oil LCO (original) white oil (50 wt %-50 wt %) density (15.6 °C) (g/mL)a paraffins (wt %) cyclo-paraffins (wt %) aromatics (wt %) SimDis (°C)b IBPc 10 wt % 30 wt % 50 wt % 70 wt % 90 wt % FBPd
0.9360 6.8 8.9 83.8
0.8152 31.1 67.7 1.1
0.8714 20.8 40.2 39.0
128.4 225.0 255.3 288.5 326.4 373.9 430.8
210.5 237.5 256.5 272.5 312.5 395.0 490.0
148.0 233.0 255.0 277.0 318.5 381.0 483.5
a ASTM D4052. b ASTM D2887. c IBP ) initial boiling point. d FBP ) final boiling point.
constant at the middle of the cell by a DP cell. The two-point thermocouple had two measurement points; one was positioned at the top of the cell to measure the vapor-phase temperature, and the other was positioned at the bottom of the cell to measure the liquid-phase temperature. The vapor phase from the top of the cell went into a low-temperature, low-pressure phase separator where the hydrocarbons in the vapor were condensed into the corresponding liquid while the uncondensed hydrogen went to a dry test meter (DTM). A stream of the vent gas after the DTM was sent to an online gas chromatograph for composition analysis. The analysis results showed negligible amounts of light hydrocarbons, indicating no significant thermal cracking during the experiments. The liquid phase from the bottom of the cell entered another low-temperature, low-pressure phase separator where the hydrocarbons were condensed and the dissolved hydrogen was released into a high-precision gas meter (Ruska 2331-801, Technel Engineering Inc.) for accurate volume measurement. The condensed liquid samples were collected, weighed, and analyzed separately for density and simulated distillation (SimDis). An overall mass balance calculation was performed to ensure that the experiments were properly conducted. The middle distillates used in this study were light cycle oil (LCO) from a commercial fluid catalytic cracking (FCC) unit, white mineral oil, and a 50 wt %-50 wt % blend of LCO and white mineral oil; these samples represent three typical types of middle distillates: highly aromatic, highly paraffinic, and something in between. Table 1 lists the key physical properties and chemical compositions of the three feeds. As shown in this table, the LCO has the highest content of aromatics, 83.8 wt % (the balance being paraffins), whereas the white oil contains only 1.1 wt % aromatics. The temperatures and pressures used in this study ranged from 250 to 400 °C and from 50 to 100 atm, respectively. The liquid feed and gas flow rates were fixed at 10 g/h and 10 NL/h, respectively, resulting in a gas-to-oil ratio of 1000 NL/kg. All of these conditions were also used in our previous hydrotreating studies with the same LCO.14-16 These experimental conditions and the types of feedstocks were carefully chosen to cover a sufficiently wide range of the conditions used either commercially or in research for hydrotreating various petroleum middle distillate fractions. The interaction coefficients obtained with the data generated from these experiments are believed to be applicable to most petroleum middle distillate hydrotreating systems, as discussed in later sections of this article. It should be pointed out that, before the experiments, VLE measurements were conducted with the toluene-hydrogen
Figure 1. Weight percentage of oil in the vapor phase at different temperatures and pressures.
system, and the data compared well with published literature data obtained under similar conditions, indicating high reliability of the experimental setup. During the experiments, two backto-back mass balance runs were conducted for some selected conditions to ensure that reproducible data were generated. In addition, experiments were repeated if any data point showed an unexpected trend compared to the other data points. 2.2. Experimental Observations. One of the important objectives of this study was to investigate the relative volatility of oils of differing compositions under different operating conditions. Figure 1a,b shows the weight percentages of oil in the vapor phase at different temperatures and different pressures, respectively, for the three oils investigated in this study. As shown in this figure, the weight percentages of the individual oils in the vapor phase increase significantly with temperature and decrease with pressure. At 250 °C (P ) 70 atm), only about 10 wt % of the oil (regardless of composition) is vaporized. In this case, VLE does not have a significant effect on the flow dynamics in the reactor. However, at 400 °C, over 80 wt % of the LCO and 100% of the white oil are found in the vapor phase, which would dramatically affect the flow dynamics in the reactor. In the pressure range of 50-100 atm investigated in this study, the weight percentage of oil in the vapor phase changed from ∼60 to ∼40 wt % at 350 °C. It is interesting to note that, for the range of operating conditions used and the same temperature and pressure, the weight percentage of oil in the vapor phase is always the highest for the white oil (except for the lowest temperature of 250 °C).
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The weight percentage is always the lowest for the LCO, and as expected, the white-oil-LCO mixture is always in the middle. This observation indicates that the volatilities of the three feedstocks are in the order white oil > mixture of white oil and LCO > LCO. From the physical and chemical composition data listed in Table 1, LCO has the highest density but much lower initial boiling point (IBP) and final boiling point (FBP), whereas the white oil has the lowest density but much higher IBP and FBP. The chemical composition (hydrocarbon type distribution) controls the relative volatility of the oil feedstocks. Therefore, it can be concluded that, at a similar boiling point, paraffins vaporize more easily than aromatics. 3. VLE Calculations and Evaluation of Interaction Coefficients 3.1. VLE Calculations during Petroleum Hydroprocessing. As mentioned earlier, to accurately model a hydroprocessing trickle-bed reactor, VLE has to be taken into account, and the chemical composition of each phase critically affects VLE. Therefore, in principle, flash calculations are required at each position at which the reactor model differential equations are integrated to calculate the temperature and conversions of individual reactants. With petroleum systems, flash calculations are normally performed by solving the equation of state (EOS). The most frequently used EOS is the Peng-Robinson equation17 P)
a(T) RT V - b V2 + 2bV - b2
(1)
where a and b are equation parameters associated with critical properties a(T) )
ΩaR2Tc2 R(T) Pc
b)
ΩbRTc Pc
(2)
For a mixture, such as petroleum fractions, mixing rules are applied to a and b N
b)
∑yb
i i
i)1
N
a)
N
∑∑yya
i j ij
(3)
i)1 i)1
aij ) (aiiajj)1⁄2(1 - dij)
(4)
In eq 4, dij is called the interaction coefficient between components i and j. For a petroleum system, it is impossible to identify all of the hydrocarbon components. Boiling-point pseudocomponents are routinely used in VLE calculations instead. The interaction coefficient dij between components i and j is critical to flash calculation. Relatively reliable correlations are available for estimating dij between hydrocarbon components. However, no such correlations are available for estimating dij between hydrogen or sour gases (H2S and NH3) and hydrocarbons, which are required for flash calculation in petroleum hydroprocessing. The evaluation of the interaction coefficients between hydrogen and hydrocarbon pseudocomponents was performed in this study by fitting the EOSsi.e., eqs 1-4sto experimental data. These data included the weight percentage of oil in the vapor phase and the SimDis data of the liquid samples from both the vapor and liquid phases from the VLE cell. 3.2. Evaluation of Interaction Coefficients. To evaluate the interaction coefficients between hydrogen and the individual hydrocarbon pseudocomponents, a commercial flash calculation program (Computer Modeling Group Ltd., denoted CMG) was
Figure 2. Computation diagram for evaluating interaction coefficients.
modified to perform the flash calculations. The flow diagram in Figure 2 shows the logic followed in the calculations. The feed oil was divided into 29 pseudocomponents according to the SimDis data, with each component covered a certain boiling-point range. Hydrogen was treated as a separate component. To use the program for flash calculation, the density, molecular weight, and boiling point of each component have to be known. All other properties, such as critical properties, are estimated by the CMG software based on these three properties. For hydrogen, all of the data are available in the open literature. To obtain the molecular weights and densities of the 29 hydrocarbon pseudocomponents, distillations were performed to generate eight fractions for each of the three feedstocks. Molecular weight, SimDis, and density were then measured for each fraction, resulting in separate correlations between molecular weight and boiling point and between density and boiling point for each feedstock. These correlations were subsequently used to estimate the molecular weights and densities of the 29 pseudocomponents. For each feedstock, a set of initial values of the interaction coefficients between hydrogen and the pseudocomponents was assumed. Flash calculations were then performed under all the experimental VLE conditions to calculate the vapor and liquid yields (weight percentage of oil in vapor) and the SimDis of the vapor and liquid products of the flash calculations. These were compared with the experimental data. Minimum discrepancies between the calculated and measured data were achieved by adjusting the values of the interaction coefficients. The interaction coefficients were assumed to vary linearly with boiling point, and the coefficients A and B in the formula below were the outcome of the parameter estimation calculations diH ) A + B × BPi
(5)
where diH represents the interaction coefficient between hydrogen and pseudocomponent i and BPi represents the boiling point of pseudocomponent i. Typical matches between the experimental and calculated data are shown in Figures 3-5. Figure 3a-c shows the calculated and measured weight percentages of oil in the vapor phase at different temperatures for LCO, white oil, and the mixture of LCO and white oil, respectively. Figure 4a-c shows the same quantities at different pressures. Figure 5a-c shows the calculated and measured SimDis data for the top and bottom liquid products for the three feeds. As shown in these figures, the calculated weight percentages of oil in the vapor phase and SimDis data match the experimental data reasonably well, except for some deviations observed for white oil (Figure 3b) at the temperatures where the oil is almost completely vaporized. This
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Figure 4. Calculated and measured weight percentages of oil in the vapor phase at different pressures.
Figure 3. Calculated and measured weight percentages of oil in the vapor phase at different temperatures.
is where the experimental measurements would be expected to be much less accurate. It is worth noting that the flash calculations predict correct shapes of the SimDis curves (Figure 5a-c), which are quite different among the three feeds, indicating a high predictive capability of the flash program if the right interaction coefficients between hydrogen and hydro-
carbons are used. It was found that the sensitivity of the interaction coefficients to the experimental data was low, and therefore, propagation of experimental errors (if any) would be insignificant. The interaction coefficients for the three feeds can be further correlated to their chemical compositions (hydrocarbon type distributions). As discussed earlier, the high aromatics content in LCO makes it less volatile than white oil, which contains mainly paraffins (see Figure 1a,b). Higher interaction coefficients for LCO were observed, suggesting a stronger interaction
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The R2 values for the linear correlations are 0.9969 for parameter A and 0.9998 for parameter B. Considering the modest sensitivity of VLE to the interaction coefficients, the simple preliminary correlations for A and B (eqs 6 and 7) are expected to be sufficiently accurate to produce reasonable flash calculation results for any middle distillate feedstock. It should be mentioned that to perform the flash calculations, the SimDis data, density, and molecular weight distributions with boiling points must be known prior to the flash calculations. SimDis analysis is easy to conduct and can be done in any petroleum laboratory. However, density and molecular weight distributions with boiling points normally require actual distillation followed by the measurements of the density and molecular weight for the individual fractions, which are expensive and time-consuming. We have developed a methodology to estimate the molecular weight distribution by boiling point using GC/FIMS (gas chromatography/field ionization mass spectrometry). The idea is to use the scanned mass of hydrocarbon ions and their abundance to calculate the molecular weight at different GC retention times, from which the boiling point of the ions can be calculated by using pre-established calibration data between retention time and boiling point. The details on this method development and some preliminary results have been presented elsewhere18 and are not discussed here. 3.3. Discussion. The VLE experiments conducted in this study were aimed at estimating the interaction coefficients between hydrogen and hydrocarbon pseudocomponents, as well as correlating the interaction coefficients with the boiling point and aromatics content of the feed. With the established correlations, VLE flash calculations can be performed at any conditions and with any type of feed. In a trickle-bed hydrotreater, the physical properties and chemical composition of the liquid (oil fraction) vary slightly (and sometimes substantially depending on hydrotreating severity) along the reactor with the progress of various hydrotreating reactions. This variation in physical properties and chemical composition will affect the VLE in the hydrotreater. One VLE flash calculation can represent only a particular position in the hydrotreater corresponding to the operating conditions and phase properties. Therefore, to accurately simulate the hydrotreater, one has to combine this VLE calculation with a reactor model to perform VLE flash calculations point-by-point, or at least within a very short axial/radial distance within which the properties and compositions of each phase can be reasonably assumed constant. With the integration of differential equations of the reactor model, the VLE flash calculations are performed many times from the inlet to the outlet of the reactor. This experimental approach is not recommended for investigating the VLE of the whole hydrotreater. Instead, it should be used to obtain required flash calculation parameters that can subsequently be used to study the VLE at any location inside the hydrotreater. To obtain an approximate estimation of VLE in a hydrotreater, it is suggested that the average physical properties of the liquid over the inlet and outlet of the hydrotreater be used to perform the VLE calculations.
Figure 5. Calculated and measured SimDis.
between hydrogen and the oil. Therefore, in order to perform flash calculations with any type of middle distillate feedstocks, the parameters A and B in eq 5 were correlated with the aromatics content (weight percentage) in the feed. Two linear correlations were obtained A ) 0.3852 - 0.00241Caromatics %
(6)
B ) 0.002245 + 1.96 × 10-5Caromatics %
(7)
4. Mapping of Hydroprocessing Trickle-Bed Reactor Operating Regimes 4.1. Calculation Basis. Taking into account the effects of VLE, in this study, the calculation and determination of operating regimes in hydrotreating trickle-bed reactors mainly focuses on an in-house pilot-plant hydrotreater. Such a hydrotreater is frequently used in hydrotreating catalyst development, evaluation, and activity stability studies to generate reliable
Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1101
LGO
LCO
an implicit assumption that conversion is less than 90% was applied. Gierman21 considered that Mears’ criterion was too conservative and, therefore, derived the relaxed criterion
0.8392 73.2 26.8
0.9360 15.7 83.8
LB 8n 1 > ln dpe PeL (1 - x)
126.5 214.5 262.5 295.0 326.5 370.5 435.0
128.4 225.0 255.3 288.5 326.4 373.9 430.8
Equation 10 was based on a 10% deviation from plug flow. In our studies on ultra-low-sulfur diesel (ULSD), we found that meeting this criterion might still be too strict because of the very high conversions (99.5% or even higher). Accordingly, an even more relaxed criterion with 15% deviation from plug flow was used in our computations
Table 2. Physical Properties and Chemical Compositions of the Feeds Used in the Computation of Flow Maps density (15.6 °C) (g/mL) saturates (wt %) aromatics (wt %) SimDis (°C)b IBP 10 wt % 30 wt % 50 wt % 70 wt % 90 wt % FBP
a
a
ASTM D4052. b ASTM D2887.
data for commercial scale-up. The specifications of the pilotplant hydrotreater used in this study are as follows: reactor i.d. ) 2.54 cm, catalyst volume ) 100-200 mL, catalyst particle size (average) ) 1.3 × 5 mm, diluent particle size ) 0.20-0.4 mm, catalyst-to-diluents volumetric ratio ) 1:1. The operating conditions used were typical for commercial middle distillate hydrotreating, as follows: temperature ) 300, 320, 340, 360 and 380 °C; pressure ) 40, 50, 60, 70, and 80 bar; liquid hourly space velocity (LHSV) ) 1.5 h-1; gas/oil ratio ) 500, 800, and 1000 NL/kg. Two typical middle distillate feeds were used; one was straight-run light gas oil (SR-LGO), and the other was the LCO used in the VLE experimental studies discussed earlier. The properties of the two feeds are listed in Table 2. As can be seen in Table 2, the two feeds are quite close to each other in terms of boiling-point range. However, there are significant differences in density and chemical composition. The LGO feed has a lower density and lower aromatics content, whereas the LCO feed has a higher density and higher aromatics content. To simplify the calculations and discussion, it is assumed that the physical properties and chemical compositions of the feeds are constant in the pilot-plant hydrotreater. 4.2. Criteria for Mapping of Operating Regimes. In trickle-bed pilot-plant hydrotreating studies, to generate reliable and consistent experimental data, it is critical to maintain the reactor operation in the desired regime, which includes plug flow of the liquid phase, complete catalyst wetting, and absence of reactor wall effects. Many published works have discussed these important issues. A separate review article is being prepared to summarize these studies. 4.2.1. Criteria for Plug Flow. In a pilot-plant trickle-bed reactor, attention should always be paid to the axial dispersion, or back-mixing in the liquid phase, especially in a short catalyst bed with a high conversion. Otherwise, the catalyst evaluation results and subsequent kinetics analysis, normally based on the assumption of plug flow, could cause deviations in reactor scaleup. An empirical criterion frequently used to verify plug flow has the simplest form19 LB > 100 dpe
(8)
Mears derived a criterion from the perturbation solution of an axial dispersion model with N mixers in series (9)
which was based on the assumption that the minimum bed length for freedom of axial diffusion should have less than 5% deviation from the bed required for plug flow. In this equation,
(11)
It should be pointed out that the Peclet number, PeL in eq 11 is based on the catalyst particle diameter and defined as uLdpe/ DL In this work, we used Hochman and Effron’s correlation22 to calculate PeL because it is one of the most frequently used in the literature. In addition, this correlation tends to give a relatively low value of PeL and, thus, a more conservative plugflow criterion. The correlation is expressed as PeL ) 0.034ReL0.53
(12)
in which the particle Reynolds number, ReL, is defined as ReL )
dpeuLFL µL
(13)
Gierman21 reported that, in radioactive tracer experiments, the particle Peclet number tended to be constant at 0.04 when the Reynolds number was less than 1.0 without a sound reason. In our opinion, there is no reason to assume a constant Peclet number even when the Reynolds number is less than 1.0. Therefore we used Hochman and Effron’s correlation (eq 12) throughout for calculations of the Peclet number. 4.2.2. Criteria for Full Catalyst Wetting. In a trickle-bed reactor, it is desirable to have all of the catalyst particles in the reactor completely covered by the flowing liquid, which gives maximum catalyst utilization. Any partial catalyst wetting could cause bypassing of the feed and, therefore, reduce the catalyst performance. If the liquid superficial velocities are very low, which is common in a pilot-plant reactor because of the relatively small volume of catalyst, partial catalyst wetting can occur. Theoretically, if the liquid flow on catalyst particle surface is dominated by the friction force rather than gravity, the fluid will tend to spread over the catalysts. On the basis of a comparison of the friction force and gravity using a large volume of experimental data, Gierman21 proposed the following criterion for wetting efficiency µLuL FLdpe2g
20
LB 20n 1 > ln dpe PeL (1 - x)
LB √20n 1 > ln dpe PeL (1 - x)
(10)
> 5.0 × 10-6
(14)
This criterion has been recommended and used by a number of researchers.23,24 4.2.3. Criteria for Absence of Reactor Wall Effects. Normally, reactor wall effects on reactor performance are not a concern for a commercial hydrotreating reactor, because the reactor diameter is much larger than the catalyst particle size. However, in a pilot-plant reactor, a concern arises when a smalldiameter reactor and relatively large catalyst particles are used. To keep the wall effects at a minimum, a trickle-bed reactor should meet the following criterion25
1102 Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009
DR > 25 dpe
(15)
This is the most rigorous criterion for wall effects reported in the literature. In our study, the calculated value was DR/dpe ) 58.2 . 25. Therefore, the reactor wall effects was negligible and no further discussion is given in this article. 4.3. Calculation Results and Discussion. 4.3.1. VLE Calculations. By using the flash calculation program with the evaluated hydrogen-hydrocarbon interaction coefficients, we calculated weight percentages of oil in the vapor phase in the pilot-plant hydrotreator at different temperatures and pressures, as shown in Figure 6a,b for the LGO and LCO feeds, respectively. The liquid hourly space velocity (LHSV) was 1.5 h-1, and the gas/oil ratio was 800 NL/kg. As shown in this figure, the predicted weight percentage of the oil in the vapor phase increased significantly with increasing temperature and decreasing pressure. For example, at a temperature of 300 °C and a pressure of 60 atm, the weight percentage of oil in the vapor phase was about 20%, representing a small but noticeable effect of VLE on the flow dynamics in the reactor. At 380 °C and 60 atm, over 60 wt % of the oil appeared in the vapor phase, indicating a significant effect of VLE on the flow dynamics in the reactor. The LCO feed showed slightly lower volatility than the LGO feed, as the former had a higher aromatic content. 4.3.2. Mapping of the Plug-Flow Operating Regime. The calculated operating regimes that meet the plug-flow criterion are shown in Figure 7a,b for the LGO and LCO feeds,
Figure 7. Mapping of plug-flow operating regime, temperature-pressure-gas/ oil ratio.
Figure 6. Calculated weight percentage of oil in the vapor phase at different temperatures and pressures.
respectively. Each curve in the figure represents, for a given gas/oil ratio, the boundary of pressure and temperature at which plug flow is achieved. The zone above the curve represents the operating temperatures and pressures that meet the plug-flow criterion, and the zone below the curve represents the operating temperatures and pressures that cannot meet this criterion. Normally, for a fixed gas/oil ratio, an increased reactor temperature requires increased pressure in order to ensure plug flow of the liquid phase. For the same temperature, an increased gas/ oil ratio also requires increased pressure to maintain liquid plug flow. Some differences were observed between the LGO feed (Figure 7a) and the LCO feed (Figure 7b). Generally, for similar reactor temperatures and gas/oil ratios, the LCO feed required a higher pressure than the LGO feed to maintain liquid-phase plug flow. This was even more pronounced at lower temperatures (below 340 °C). As seen in Figure 7b, in the lowtemperature range, there was even an inverse trend that reduced the temperature at which an increase in pressure was required to maintain plug flow. This was due to the significant increase in viscosity for the LCO feed with decreasing temperature. To maintain the Reynolds and Peclet numbers, higher liquid velocity/flow rate was needed, which required higher pressure to increase the amount of oil in the liquid phase. This important finding indicates that, even for similar middle distillates, such as LCO and LGO with similar boiling-point ranges, different liquid densities and chemical compositions can result in different flow dynamics and, therefore, reactor performance.
Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1103
Figure 8. Mapping of plug-flow operating regime, temperature-pressurecatalyst bed length.
Figure 9. Mapping of plug-flow operating regime, temperature-diluent particle size-pressure.
For some pilot-plant studies, the operating conditions (temperature, pressure, LHSV, and gas/oil ratio), catalyst particle size, and diluent particle size are fixed according to process or test requirements. Therefore, to maintain plug flow of the liquid inside the reactor, the catalyst bed length (or catalyst volume) has to meet a minimum value under a particular set of operating conditions, which can be clearly seen in eq 11. For a given temperature, a minimum catalyst bed length can be calculated while maintaining other operating parameters constant by applying eq 11. Figure 8a,b shows the requirement for minimum catalyst bed length at different reactor temperatures and pressures for a fixed LHSV of 1.5 h-1 and a gas/oil ratio of 800 NL/kg for the LGO and LCO feeds, respectively. Each curve in the figure represents, for a given pressure, the boundary of temperature and catalyst bed length at which plug flow is met. Again, for each curve in the figure, the zone above the curve represents the operating temperatures and catalyst bed lengths that meet the plug-flow criterion, and the zone below the curve represents the operating temperatures and catalyst bed lengths that cannot meet the plug-flow criterion. As seen in this figure, an increased temperature requires an increased catalyst bed length (or catalyst volume) to achieve plug flow of the liquid. At higher temperatures and lower pressures, the required catalyst bed length increased sharply. For similar operating conditions, the LCO feed (Figure 8b) required longer catalyst beds (or greater catalyst volumes) than the LGO feed (Figure 8a) to maintain plug flow of the liquid in the reactor.
To achieve a better liquid and gas distribution in the catalyst bed and to have better heat transfer, it is a common practice to use smaller diluent particles in pilot-plant reactors. Using smaller diluent particles also improves the flow dynamics and helps in meeting the plug-flow criterion. Therefore, if the operating conditions, catalyst size, and volume are fixed, changing the diluent particle size (with a fixed catalyst-to-diluent ratio of 1:1) can also achieve plug flow of the liquid phase. Figure 9a,b shows the requirement for diluent particle size at different reactor temperatures and pressures for a fixed LHSV of 1.5 h-1 and gas/oil ratio of 800 NL/kg for the LGO and LCO feeds, respectively. Each curve in the figure represents, for a certain pressure, the boundary of temperature and diluent particle size at which plug flow is met. For each curve in the figure, the zone below the curve represents the operating temperatures and diluent particle sizes that meet the plug-flow criterion, and the zone above the curve represents the operating temperatures and diluent particle sizes that cannot meet the plug-flow criterion. In this figure, an increase in temperature requires a reduced diluent particle size. For similar operating conditions, the LCO feed (Figure 10b) requires smaller diluent particles than the LGO feed to maintain plug flow of the liquid in the reactor. 4.3.3. Mapping of the Full Catalyst Wetting Regime. As discussed earlier, the VLE in the pilot-plant hydrotreater could significantly change the liquid-phase flow rate as a result of partial vaporization of the oil into the vapor phase. Such a change in liquid flow rate could result in incomplete catalyst wetting, leading to bypassing of the feed and, therefore, reduced
1104 Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009
Figure 11. Mapping of full catalyst wetting operating regime, temperaturediluent particle size-pressure.
Figure 10. Mapping of full catalyst wetting operating regime, temperaturepressure-gas/oil ratio.
catalyst performance. The extent of partial catalyst wetting is normally random and differs from one set of conditions to another. Therefore, it is important to maintain full/complete catalyst wetting in a pilot-plant hydrotreater in order to generate reliable and reproducible data. Unlike the plug-flow criterion, the criterion for full catalyst wetting is not related to the catalyst bed length. It is related only to the operating conditions and the catalyst and diluent particle sizes. Figure 10a,b shows the mapping of full catalyst wetting for the LGO and LCO feeds, respectively. Each curve in the figure represents, for a certain gas/oil ratio, the boundary of pressure and temperature at which full catalyst wetting is achieved. The zone above the curve represents the operating temperatures and pressures that meet the full catalyst wetting criterion, and the zone below the curve represents the operating temperatures and pressures that cannot achieve full catalyst wetting. Normally, for a fixed gas/oil ratio, an increased reactor temperature requires increased pressure to ensure full catalyst wetting. For the same temperature, an increased gas/oil ratio also requires increased pressure to maintain full catalyst wetting. It is interesting to see in Figure 10a,b that, for the same temperature and gas/oil ratio, the LGO and LCO feeds require similar pressures to have complete catalyst wetting, which is different from the plug-flow mapping discussed earlier where the LCO feed required a higher pressure than the LGO feed to maintain plug flow in the liquid at the same temperature and gas/oil ratio (see Figure 7a,b). By comparing Figures 7a,b and 10a,b, it can be seen that the full catalyst wetting criterion is easier to meet, or more relaxed, than the plug-flow criterion.
Therefore, as long as the plug-flow criterion is met, full catalyst wetting will automatically be achieved for the systems considered in this study. Similar calculations were performed to determine the maximum diluent particle size to meet the full catalyst wetting criterion. Figure 11a,b shows the maximum diluent particle size at different reactor temperatures and pressures for a fixed LHSV of 1.5 h-1 and gas/oil ratio of 800 NL/kg for the LGO and LCO feeds, respectively. Each curve in the figure represents, for a given pressure, the boundary of diluent particle size and temperature at which full catalyst wetting is achieved. The zone below the curve represents the operating temperatures and diluent particle sizes that meet the full catalyst wetting criterion, and the zone above the curve represents the operating temperatures and diluent particle sizes that cannot achieve full catalyst wetting. Increased temperature and decreased pressure require reduced diluent particle size. However, the two feeds did not show much difference in the trend under the same conditions, with the LCO feed requiring slightly lower diluent particle sizes. This might be explained with the full wetting criterion (eq 14), which is inversely proportional to the square of the effective particle size. In eq 14, the effective particle size (dpe) has more weight in the magnitude of the left-hand side of the equation than the other terms (velocity, density, and viscosity). This might also explain why operating regimes satisfying the full catalyst wetting criterion are more relaxed than those determined by the plug-flow criterion, as the diluent particle size (0.25 mm) used in the calculation was so small that full catalyst wetting would not become a problem. If a larger diluent particle size were chosen for the calculation, the full catalyst wetting criterion might supersede the plug-flow criterion.
Ind. Eng. Chem. Res., Vol. 48, No. 3, 2009 1105
4.3.4. Further Discussion. The results and discussion provided in this article are related to pilot-plant hydrotreating reactors and cannot be directly extended to commercial installations. In a commercial hydrotreating reactor that is operated at a similar LHSV and gas/oil ratio, the liquid and gas velocities might be significantly different from those in the pilot-plant reactor because of the differences in catalyst bed length and diameter. Generally speaking, LB/dpe in a commercial hydrotreating reactor is much larger than that in a pilot-plant reactor, so the plug-flow criterion is easier to meet in the commercial reactor. However, because a commercial hydrotreating reactor is normally operated in adiabatic (or close to adiabatic) mode, there is a significant temperature gradient in the axial direction (and possibly in the radial direction as well). Such a temperature gradient would make the vapor-liquid equilibrium inside the reactor much more complicated, which would subsequently affect the establishment of plug flow and full catalyst wetting. Therefore, to ensure plug flow and full catalyst wetting in a commercial hydrotreating reactor, the maximum possible temperature in the reactor should be used in the calculation. Detailed calculation results and discussion on this topic will be reported in future publications. 5. Conclusions In this work, VLE experiments were conducted in a continuous-flow unit with hydrogen and petroleum middle distillates, including an LCO from an FCC process, white oil, and a mixture of LCO and white oil, under typical hydrotreating conditions. The interaction coefficients between hydrogen and hydrocarbon boiling-point pseudocomponents were evaluated by fitting the calculated VLE data with the experimental data. The calculated VLE results were further used to predict the flow dynamics in pilot-plant hydrotreaters with two different middle distillate feeds (SR-LGO and LCO). Further calculations were performed to provide a mapping of operating conditions under which the desired operating regimes (plug flow, full catalyst wetting, absence of reactor wall effects) could be maintained. The following conclusions can be drawn from this study: (1) The relative amount of oil in the vapor phase increases significantly with increasing temperature and decreasing pressure. (2) The volatilities of the feedstocks are affected by, in addition to the boiling point and molecular weight, the aromatics content in the oil feed: the higher the aromatic content, the lower the volatility. (3) The correlated hydrogen-hydrocarbon interaction coefficients could reasonably simulate flashing of any type of feedstock if the SimDis, density, molecular weight distributions with boiling point, and aromatic contents are known. (4) The calculation of pilot-plant hydrotreater operating regimes showed that hydrotreating the LCO feed requires a higher pressure, a lower gas/oil ratio, a longer catalyst bed, or smaller diluent particles than hydrotreating the LGO feed in order to maintain plug flow of the liquid in the reactor. However, hydrotreating the two feeds requires similar operating parameters to achieve full catalyst wetting. Acknowledgment The authors are grateful to the following individuals for their help and support: Dennis Carson for conducting the VLE experiments. Partial funding for NCUT has been provided by the Canadian Program for Energy Research and Development (PERD), the Alberta Research Council (ARC) and the Alberta Energy Research Institute (AERI). The authors are grateful to
NCUT’s staff from the pilot plant and analytical laboratory for their support and help. The editing of this manuscript by Norman Sacuta is greatly appreciated. Nomenclature A ) fitted parameter used to determine interaction coefficients a ) parameter of the Peng-Robinson equation B ) fitted parameter used to determine interaction coefficients b ) parameter of the Peng-Robinson equation BP ) boiling point (°C) Caromatics % ) aromatic content of feed (wt %) dij ) interaction coefficient between components i and j diH ) interaction coefficient between pseudocomponent i and hydrogen dpe ) effective particle size (cm) DL ) liquid diffusion coefficient (cm2/s) DR ) diameter of reactor bed (cm) g ) acceleration of gravity (981 cm/s2) LB ) reactor bed length (cm) LHSV ) liquid hourly space velocity (h-1) n ) reaction order P ) pressure (atm) Pc ) critical pressure (atm) PeL ) particle Peclet number, uLdpe/DL R ) gas constant (83.14472 cm3 atm K-1 mol-1) ReL ) particle Reynolds number T ) temperature (°C or K) Tc ) critical temperature (K) uL ) liquid superficial velocity (cm/s) V ) volume of gas (cm3) x ) conversion of component y ) fraction of component Greek Letters µL ) liquid-phase viscosity (cP) FL ) liquid-phase density (g/cm3) Ωa ) constant (0.45724) Ωb ) constant (0.07780) Subscripts i ) component index j ) component index L ) liquid phase
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ReceiVed for reView April 14, 2008 ReVised manuscript receiVed May 27, 2008 Accepted July 8, 2008 IE8006006
