Coke Deposition Profiles during Artificial Aging of Hydroprocessing

In order to enhance coke deposition, experiments were done at high temperature (573−673 K), low ... Abdulazem Marafi , Antony Stanislaus , Edward Fu...
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Ind. Eng. Chem. Res. 2007, 46, 421-429

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Coke Deposition Profiles during Artificial Aging of Hydroprocessing Catalysts Bas M. Vogelaar,*,† Jeroen Gast, Erwin M. Douma, A. Dick van Langeveld, Sonja Eijsbouts,‡ and Jacob A. Moulijn Delft UniVersity of Technology, Faculty of Applied Sciences, Department of Reactor and Catalysis Engineering, Julianalaan 136, 2628 BL Delft, The Netherlands

A model NiMo/Al2O3 catalyst was aged under various process conditions during the hydroprocessing of light gas oil. In order to enhance coke deposition, experiments were done at high temperature (573-673 K), low pressure (1-5 MPa) and hydrogen to oil (HTO) ratio (190-560 Nm3H2/m3oil), and with the addition of polyaromatic compounds. No catalyst deactivation was observed during the initial 50 h on stream. At high HTO ratio a moderate coke deposition was observed, almost independent of pressure and temperature. At low HTO ratio, the conversion decreased and the amount of coke increased, due to hydrogen depletion. At low HTO ratio, the addition of polyaromatics lowered the HDS conversion, which was found not to be due to inhibition but rather to competition in hydrogen consumption. Aromatics addition did not enhance coking, however. During aging, a significant fraction of the catalyst initial pore volume was blocked by coke deposits. Calculations predicted no diffusion limitations for the aromatic and sulfur-containing reactants but significant hydrogen concentration gradients inside the catalyst pellets at low pressure. This mass transfer limitation of hydrogen resulted in the formation of typical bell-shaped coke deposition profiles, which could be explained by a simple model assuming the rate of coke deposition being inversely proportional to the hydrogen partial pressure. Introduction The deactivation behavior of a catalyst is extremely important in the design of catalytic processes; it determines the time of stable operation before the catalyst needs to be replaced or regenerated. One of the important reasons why a hydroprocessing catalyst deactivates is, like for many other heterogeneous catalysts, coke deposition. Coke may selectively deposit on the active sites thereby lowering the intrinsic activity or plug the pores of the catalyst causing diffusion limitations. Understanding of this deactivation process is essential for the development of long-life catalysts.1 Especially, the rapid initial deactivation during the first few hours on stream is interesting, as the level of coke on the catalyst can easily reach 10 wt % in that period of time and is usually accompanied by a significant drop in pore volume and surface area. Studies on the factors influencing the initial coking behavior in an industrial reactor and its effect on catalyst properties are scarce, but several researchers attempted to elucidate this process by performing artificial aging experiments on lab scale and analyzing spent catalysts from industrial hydrotreaters.1-11 Their observations are similar: coke builds up fast during the first few hours of operation and then reaches a stable level. The catalysts show a considerable decrease in pore volume and surface area. The (aromatics) hydrogenation (HDA) and hydrodenitrogenation (HDN) activity is much more affected than the hydrodesulfurization activity (HDS). The literature is, however, less conclusive about the mechanism of deactivation; some authors suggest pore blocking, while others propose active site poisoning as the major cause for the activity loss. When excessive pore blocking occurs, the accessibility of the catalyst pore network is reduced, and diffusion limitations * To whom correspondence should be addressed. Tel.: +31 206347648. Fax: +31 206347653. E-mail: Bas.Vogelaar@ albemarle.com. † Current address: Albemarle Catalysts, P.O. Box 37650, 1030 BE Amsterdam, The Netherlands. ‡ Albemarle Catalysts, P.O. Box 37650, 1030 BE Amsterdam, The Netherlands.

may arise for the different reacting species, especially for large molecules. One can check the presence of mass transfer limitations by calculating the external and internal effectiveness factor for the given reaction.12 For external (film layer) mass transfer the following equations are used:

ηe ) (1 - Ca)n LrV,p kfCb

(2)

ShDb dSiC

(3)

Ca )

kf )

(1)

in which ηe is the external effectiveness factor, Ca is the Carberry number, n is the reaction order, L is the diffusion length (half of the catalyst pellet radius, in m), dSiC is the diameter of the SiC diluent particles (m) generally used in laboratory tests, rV,p is the observed volumetric reaction rate inside the pellet (mol/m3/s), kf is the transfer coefficient for the stagnant liquid film around the particles (m/s), Db is the bulk diffusivity (m2/ s), Cb is the bulk reactant concentration (mol/m3), and Sh is the Sherwood number. The Sherwood number is a function of the liquid velocity, but as this velocity is very low in the case of a laboratory reactor, the lower limit value Sh ) 2 was used. For mass transfer inside the pellet we used the following equations:

ηi )

Φ)

Φ φ2

(4)

L2rV,p n + 1 DeffCs 2

(

I1(2φ) Φ)φ I0(2φ)

)

(5)

(6)

where ηi is the internal effectiveness factor, Φ is the Weisz

10.1021/ie060870a CCC: $37.00 © 2007 American Chemical Society Published on Web 12/20/2006

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Figure 1. Normalized reactant concentration profiles for different internal effectiveness factors (indicated above each curve) for first-order (A) and secondorder reaction (B).

number, φ is the Thiele modulus, Deff is the effective diffusion coefficient (m2/s), Cs is the reactant concentration at the catalyst surface (mol/m3), and I0 and I1 are Bessel functions. This set of equations can be solved numerically when the surface concentration is known from the following:

Cs ) Cb(1 - Ca)

(7)

The solution of the above equations will yield the value for the Thiele modulus and both effectiveness factors. Hence, the presence of external and/or internal mass transfer limitations can be assessed, based on the observed reaction rates. A second, independent property can be used to check for internal diffusion limitations: the presence of coke profiles in the catalyst pellet. Nonuniform coke deposition can only occur when there are significant diffusion limitations of one or more species involved in the coking process. The concentration profile of a reactant in a cylindrical pellet can be calculated by solving the differential equation for the mass balance:

8 1 d dc φ2cn x ) x dx dx n+1

( )

(8)

in which c is the dimensionless reactant concentration C/Cs and x is the dimensionless radius r/rp. Figure 1 shows the resulting profiles for a first- and second-order reaction with various effectiveness factors. The reaction rates are highest at the inlet of the catalyst bed; hence, the effectiveness factors for the catalyst pellets at the top of the bed will be lowest. The most pronounced coke profiles are therefore likely to be found in the top section of the bed. For a first-order reaction, however, the effectiveness factor is constant at every position in the catalyst bed, because the diffusion rate is first order as well. Higher order kinetic behavior is generally observed when a wide range of compounds, which may have different reactivities, are lumped together into one kinetic equation. In this study we used this approach to describe the conversion of the total of sulfur and aromatic compounds. In practice, reaction orders between 1 and 2 are found. We investigated both limiting cases, i.e., all results were evaluated using both first- and second-order kinetics. In this study we want to investigate both intrinsic deactivation (poisoning) and pore blocking (fouling) by monitoring the activity and the diffusion properties of the catalyst during the initial period of operation in the hydroprocessing of light gas oil (LGO). The occurrence of mass transfer limitations during the coking process was studied by analyzing the coke profiles in the spent catalyst particles. Mass transfer properties of the reactants were calculated using (estimated) diffusivity values

Table 1. Properties of the NiMo Catalyst Used in This Studya property

value

method

MoO3 content NiO content surface area (BET) pore volume mean pore diameter particle density porosity tortuosity

12 wt % 2.5 wt % 292 m2/g 0.53 cm3/g 7.3 nm 1.3 g/cm3 0.6 1.3

N2-BET N2-BET N2-BET estimated (ref 13) estimated (ref 13) estimated (ref 13)

a

Analysis data were supplied by Albemarle Catalysts.

Table 2. Properties of the Light Gas Oil (LGO) Used in This Studya property

value

sulfur content nitrogen content mono aromatics di aromatics tri+ aromaticsb initial boiling point boiling point at 50 vol% (T50%) final boiling point density at 15 °C (288 K) average molecular weight viscosity at 40 °C (313 K)

8445 ppm 124 ppm 16.2 wt % 9.9 wt % 1.3 wt % 233 °C 316 °C 387 °C 854.2 kg/m3 0.25 kg/mol 5.3 mPa‚s

method

HPLC HPLC HPLC ASTM D86 ASTM D86 ASTM D86 eq 9 eq 12

a Analysis data were supplied by Albemarle Catalysts. b Polynuclear aromatics with three or more aromatic rings.

and kinetic properties obtained from the above equations. Calculations in this work were made for both the first-order and the second-order case and at inlet conditions. It is generally believed that coke formation starts from large polyaromatic structures;13 therefore, in some experiments a mixture of 1 wt % naphthalene, 0.5 wt % anthracene, and 0.5 wt % pyrene was added as coke precursor. In order to further accelerate coke deposition, some of the artificial aging experiments were performed at high temperature, low pressure, and low hydrogen to oil ratio. Experimental Section We used a NiMo-based model catalyst supplied by Albemarle Catalysts, which is further specified in Table 1. The catalyst was supported on a high-purity γ-Al2O3 and did not contain phosphorus, to rule out possible coking effects by phosphate surface groups or contaminations in the support material. The particle density, porosity, and tortuosity were estimated from literature data on comparable catalysts.14 As a feedstock we used an LGO, also supplied by Albemarle Catalysts, of which the details are given in Table 2. Additional properties like (average) molecular weight, density, and viscosity at reaction conditions were estimated by numerical

Ind. Eng. Chem. Res., Vol. 46, No. 2, 2007 423 Table 3. Diffusional Properties of Reactants at Process Conditions property

anthracene

naphthalene

hydrogen

molar volume Vm (cm3/mol) bulk diffusivity at 573 K (m2/s) bulk diffusivity at 623 K (m2/s) bulk diffusivity at 673 K (m2/s) hydrodynamic diameter dh (nm) effective diffusivity at 573 K (m2/s) effective diffusivity at 623 K (m2/s) effective diffusivity at 673 K (m2/s) ref

214 2.0 × 10-8 4.3 × 10-8 9.5 × 10-8 0.85 0.59 × 10-8 1.3 × 10-8 2.9 × 10-8 13

157 2.4 × 10-8 5.2 × 10-8 11 × 10-8 0.74 0.71 × 10-8 1.6 × 10-8 3.4 × 10-8 13

14.3 10 × 10-8 22 × 10-8 48 × 10-8 0.30 3.9 × 10-8 8.6 × 10-8 19 × 10-8 11

fitting and extrapolation of published data for a number of oil fractions,15,16 using the following:

M ) 7.25 × 10

-7

T50%2

(9)

F ) F0 - 0.723(T - T0)

(10)

η ) η0 exp[-0.0562M(T - T0)]

(11)

η40° ) 1.09 × 10-4 exp(15.5M)

(12)

where M is the average molecular weight (kg/mol), T50% is the temperature at 50% of the boiling curve (K), F is the density (kg/m3), F0 is the density at reference temperature T0, η is the viscosity (Pa‚s), η0 is the viscosity at reference temperature T0, and η40° is the (reference) viscosity at 40 °C (313 K) in Pa‚s. According to the observations of Boned et al.15 the effect of pressure on these properties in the range of 0-5 MPa is negligible. The effective diffusivities, required for the mass transfer calculations, were calculated according to the following:14

 Deff ) Dbulk F(λ) τ

(13)

in which Deff is the effective diffusion coefficient (m2/s), Dbulk is the bulk diffusivity (m2/s),  and τ are the porosity and

tortuosity of the catalyst, and F(λ) is the restrictive factor due to hindered diffusion. This factor can be approximated using the following:14

( )

F(λ) ) (1 - λ)z ) 1 -

dh dp

4

(14)

where λ is the ratio between the average pore diameter dp (nm) and the hydrodynamic diameter of the solute dh (nm) and z ) 4 for λ < 0.2. The bulk diffusivity Dbulk (m2/s) was estimated using the Wilke-Chang equation:12

Ms0.5T Dbulk ) 5.9 × 10-17 ηsVm0.6

(15)

in which Ms is the molecular mass of the solvent (kg/mol), ηs is the viscosity of the solvent (Pa‚s) at temperature T (K), and Vm is the molar volume of the solute at its normal boiling point (m3/mol). The sulfur compounds in gas oil that determine the HDS rate are dibenzothiophene and its substituted derivatives. The diffusional properties of these molecules were approximated using the values for anthracene. For the aromatic species, the diffusion properties of naphthalene were used. These properties, together with the values for hydrogen, are summarized in Table 3. The catalysts were tested is a microflow setup depicted in Figure 2A. Liquid was fed from a standpipe (for flow indication) using a controlled high-pressure pump, mixed with a hydrogen or H2S/H2 gas flow, and preheated to reaction temperature before it enters the reactor. It was verified, using naphthalene hydrogenation experiments under similar conditions (not shown), that the hydrogen transfer rate to the liquid was sufficient for our kinetic HDS tests. The catalyst bed was diluted with fine SiC particles (400 µm) to prevent bypassing of the catalyst. The amount of catalyst used for every test was 4.0 g, giving a total bed volume of 8.2 cm3. The rest of the reactor was filled with fine (400 µm) and coarse (200 µm) SiC particles as shown in

Figure 2. (A) Microflow apparatus for catalyst performance testing. (B) Photograph of the catalyst bed.

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Table 4. Experimental Conditions and Steady-State Conversions expt code

aging method

feed

p (MPa)

T (K)

LHSV (1/h)

HTO (nL/L)

XS (%)

XA (%)a

XH2 (%)b

S1 S2 S3 S4 A1 A2 A3 A4

standard standard standard standard accelerated accelerated accelerated accelerated

LGO LGO LGO LGO LGO LGO LGO + aromatics LGO + aromatics

5.0 1.0 5.0 1.0 5.0 1.0 5.0 1.0

573 573 623 623 673 673 673 673

1.9 1.9 3.7 3.7 6.6 6.6 6.6 6.6

510 510 510 510 560 190 560 190

40 40 73 80 76 54 74 32

40 40 73 80 76 54 74 32

9 9 18 19 17 36 17 23

a

Conversion of aromatics was assumed to be equal to the sulfur conversion (see text). b Conversion of hydrogen was calculated using eq 20.

Figure 2B. After the reactor gas and liquid were separated using a level controller with a liquid lock, the off gas was sent to the vent after neutralization of the H2S in a NaOH scrubber and the liquid was sent to a computer-controlled autosampler. The catalyst was sulfided in situ in the gas phase using 10 vol % H2S in H2 at 60 mL/min and atmospheric pressure. The temperature was held at 298 K for 1 h, raised at 3 K/min, and held at 673 K during 2 h. After that, the reactor was cooled down. To ensure proper wetting of the catalyst bed, the reactor was stabilized at room temperature during 6 h using the required liquid and gas flow rates. After that, the reactor was rapidly heated to the required temperature in 1 h. Samples were taken at 15 min to 5 h intervals and analyzed off-line using a GC with a sulfur-specific detector (Sievers SCD 355). The HDS conversion XS was calculated by integrating the total area of all sulfur compounds detected by GC and comparing it to the LGO feed. The aromatics (HDA) conversion XA could not be measured, because we had no specific analysis techniques for aromatics at our disposal. Therefore, the aromatics conversion and hydrogen consumption were estimated, as described in the discussion below. Two series of experiments were carried out: The first series were considered standard aging tests under relatively low temperatures and a high HTO ratio. These tests were labeled “SX” where X is the number of the test. The second series were accelerated aging tests labeled “AX”, which were preformed at high temperature, some at low HTO ratio and with addition of polyaromatics. The experimental conditions of each test are shown in Table 4. Coke profiles were measured using a Renishaw Ramascope System 2000 instrument. Analyses were carried out in a flow cell (Linkam THMS600) purged with dry nitrogen to protect the coked samples from oxidizing. Samples were prepared by cutting several extrudates into small pieces of about 1 mm thickness and sticking them onto a glass support disk using double-sided adhesive tape. The “coke signal” intensity was calculated by integrating the Raman bands between 950 and 1750 cm-1. The carbon content of the spent catalysts was measured using a Leco CS 225 analyzer. The sample was combusted in an O2 gas stream, and the CO2 produced was quantified with an infrared detector. The BET surface area was measured by means of a Quantachrome Autosorb 6B unit using N2 physisorption at 77 K. Prior to the analyses the samples were heated at 473 K in vacuum during 16 h. Pore volumes were measured by Hg porosimetry using a CE Instruments Pascal 140-440 apparatus. Results Surprisingly, we did not observe significant catalyst deactivation in any of the experiments. After a short run-in period, the HDS conversion was stable during the entire run (over 50 h on stream) as shown in Figure 3. One experiment was extended to 250 h, during which the activity remained stable (not shown).

Figure 3. Examples of activity curves at 5 MPa pressure (experiments S1, S3, and A1). After a short run-in time (indicated by arrows) the catalyst activity is stable.

The run-in period (indicated by arrows) is not only caused by the time to warm up the reactor but also reflects the response time of the system, as it became shorter at increasing space velocity. Figure 4A shows the HDS conversion as function of temperature and pressure. The activity of the catalyst increases with temperature, but it is striking that the pressure seems to have no influence on conversion. The coke levels in Figure 4B show that the total amount of coke is approximately equal for the standard aging tests (around 6 wt %). When we look at the accelerated aging tests shown in Figure 5A, the addition of polyaromatics did not have much influence on the HDS conversion at 5.0 MPa pressure. At 1.0 MPa and lower HTO ratio, however, polyaromatics have a huge negative impact on HDS conversion. In both cases the coke levels on the spent catalysts are virtually the same with or without aromatics addition (Figure 5B). We do, however, observe that the amount of coke doubles when the catalyst is subjected to a low pressure (1.0 MPa) and a low HTO ratio (190 m3H2/m3oil). The total pore volume and surface area of the fresh catalyst and all spent catalysts are listed in Table 5, together with the average pore diameter, the coke content, and the time on stream for each sample. Figure 6 shows the pore size distributions of the freshly sulfided and two aged catalysts (S3 and S4). The distributions for the other spent samples were similar to those of catalyst S3 and S4 and are therefore not shown. All spent catalysts show a significant decrease in pore volume and surface area, whereas the average pore diameter only decreased slightly. The effect of the coke deposits on the pore structure of the catalyst is shown in Figure 7, where the results of all of the tests are plotted in the same graph. At 10 wt % deposited coke the surface area has been reduced by approximately 25% and the pore volume by approximately 30%. The coke profiles of the spent catalysts were found to depend only on the applied pressure in the process. All catalysts aged at 5.0 MPa showed homogeneous profiles, whereas all catalysts

Ind. Eng. Chem. Res., Vol. 46, No. 2, 2007 425

Figure 4. HDS conversion and coke content for the standard aging tests. All tests were performed at an HTO of 510 nL/L; the other conditions are indicated in the graph.

Figure 5. HDS conversion and coke content for the accelerated aging tests. All tests were performed at 673 K; the other conditions are indicated in the graph. Table 5. Coke Content, Pore Volume, Surface Area, and Average Pore Diameter of the Spent Catalysts catalyst

time on stream coke content surface area pore volume dpore (cm3/g) (h) (wt %) (m2/g) (nm)

fresh S1 S2 S3 S4 A1 A2 A3 A4

248 97 144 76 30 28 30 28

0.05 5.3 6.4 5.9 5.8 4.3 8.7 4.4 8.7

282 224 214 220 228 228 213 235 216

0.54 0.40 0.38 0.40 0.41 0.45 0.39 0.45 0.38

7.3 6.8 6.7 6.9 6.8 7.5 7.0 7.3 6.7

subjected to 1.0 MPa exhibited core poisoning. Most of these samples had very typical bell-shaped profiles, where the amount of coke gradually increases from the outer surface of the particle and has a sharp maximum at the center. As an example, two profiles of the catalysts aged at 623 K (S3 and S4) are displayed in Figure 8. Discussion During the first hours on stream no deactivation could be observed in any of the experiments. This may have two trivial reasons: either the deactivation takes place during the warmup period of the reactor, which takes 1-2 h, or there is no measurable deactivation in this time window. One case has been reported in which the amount of coke reached a stable value within 2 h on stream.5 In another case it was found that at low coke levels (below 5 wt %) the deactivation is negligible.7 Some of the catalysts in our study have significantly higher amounts of coke, however. The most probable reason is the low coking sensitivity of the HDS reaction in particular, which is consistent with the literature.1-3,6,8,9 The effect of aging time on the amount of coke is unclear, as shown by the coke levels in Table 5. Catalysts processed at high pressure showed similar coke levels

Figure 6. Pore size distribution for the fresh catalyst and aged catalysts S3 and S4.

as the ones processed at low pressure, yet their time on stream was much longer. Most artificial aging studies in the literature report that the coke amount reaches a steady state within 20 h on stream2,4-6,9, with exception of Zeuthen et al.7 who observed a steady state after 80 h on stream. After this fast initial coke buildup the further accumulation of coke is usually much slower.17 Therefore, we find it reasonable to assume that the coke levels on the aged catalysts all have reached their steadystate level within 20 h on stream. The coke content appears to be almost independent of pressure as well as temperature (Figure 4). The catalysts aged at 673 K and 5 MPa had the lowest amount of coke (Figure 5). This was not expected, because in general coke deposition is accelerated at high temperature and low hydrogen pressure. However, in the case of hydroprocessing catalysts, contradicting observations are reported in the literature: Gualda and Kasztelan6 showed that the coke content declined by a factor of 0.8 as pressure increased from 2 to 5 MPa, whereas Richardson et al.4 observed no change in coke content between 7 and 10 MPa. A possible explanation could be that at the high hydrogen to oil ratio (>500) applied in our experiments, the hydrogen supply

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Figure 7. Pore volume and surface area as function of coke content for the fresh catalyst (cross) and spent catalysts after standard (dots) and accelerated (squares) aging.

Figure 8. Typical coke profiles after artificial aging at 5 MPa (S3) and 1 MPa (S4) pressure, measured by Raman spectroscopy.

rate is not limiting the coke deposition rate, even at low pressure. As for the temperature independency, Furimsky and Massoth17 reported that the amount of deposited coke was constant between 648 and 713 K. Their explanation is that, in this temperature range, the hydrogenation of coke precursors competes effectively with the coke deposition reactions. Another remarkable result is the pressure independency of the HDS conversion (Figure 4), while it is generally known that the HDS reaction rate is a strong function of hydrogen pressure. This observation can be explained by a change in H2S partial pressure at varying total pressure18 and is typical for reactions in which one of the products decreases the rate of reaction. When both pressure and HTO ratio are decreased (experiment A2), we do see an effect on HDS conversion and coke buildup. Clearly in this case not enough hydrogen is available to facilitate both HDS and HDA and suppress the coking reactions. Table 4 shows that in this case the hydrogen consumption reaches almost 40%, based on calculations given below. The problem of hydrogen availability is further demonstrated by the experiment with polyaromatics addition at the same conditions (experiment A4). A huge drop in HDS conversion is the result (Figure 5), which is due to competition between the HDS and HDA reactions for the available hydrogen. The same experiments (A1 and A3) at high pressure and HTO ratio prove that this effect is not due to inhibition of the catalyst, as often reported in the literature.19 When this would be an inhibitory effect, the addition of polyaromatics should also decrease the HDS conversion at high pressure (experiment A3 vs A1). This is clearly not the case, as can be seen in Figure 5. Finally, the added polyaromatics had no effect on the coke content at both low and high hydrogen availability. This could imply that the chosen compounds (naphthalene, anthracene, and pyrene) do not have a high affinity to form coke in comparison to the other components in the feed or that the added amount was too low to see an effect.

Apparently, for all spent catalysts the type of coke is similar, because in the plots of pore volume and surface area versus coke content all points are very close to the same trendline (Figure 7). This means that the density of the coke is the same in all cases. The pore size distribution mainly changed in intensity (Figure 6), and the mean pore diameter decreased slightly as shown in Table 5. A similar observation was reported by Wood and Gladden20 for industrial spent hydroprocessing catalysts. Analogously to their conclusion, we assume that the decrease in average pore diameter is caused by the deposition of a thin coke layer on the pore walls. This amount of coke does not lead to any significant deactivation, probably because the major part of the active phase is still accessible. This is in line with the model of Richardson et al.4 who suggest that the active sites remain free of coke because of their hydrogenation capabilities. Van Doorn and Moulijn21 envisage that the active phase protrudes through the coke layer on the support into the reaction medium. If indeed these models agree with our experiments, poisoning of the active sites does not take place under the applied conditions. However, this effect alone cannot explain the major loss of pore volume observed in Figures 6 and 7. Hence, we assume that a fraction of the pore network is completely blocked by coke. The fact that no deactivation is observed after the “run-in” period leads us to conclude that a major part of the coke deposits during the first hours on stream. Because the conversion of aromatic compounds and hydrogen could not be measured, the following assumptions were made based on general observations of the kinetic and thermodynamic properties of the HDA reaction:22-25 The hydrogenation rate of the first aromatic ring of polyaromatics (tri+) and diaromatics is much higher than that of monoaromatics. Unpublished results provided by Albemarle Catalysts show that under typical hydrotreating conditions (T ) 603 K, p ) 4 MPa, LHSV ) 2, HTO ) 300, LGO feed), the conversion of monoaromatics is about 20%, that of diaromatics

Ind. Eng. Chem. Res., Vol. 46, No. 2, 2007 427

about 90%, and the conversion of polyaromatics (tri+) is close to 100%. The resulting total aromatics conversion is around 50% under these conditions. For the complete conversion of one monoaromatic molecule, three hydrogen molecules are required, for diaromatics five, and for polyaromatics seven H2 molecules (i.e., all polyaromatics are assumed to be triaromatics). The total hydrogen consumption normalized per aromatic molecule converted (nH2,A) can now be estimated using the following:

3CmonoXmono + 5CdiXdi + 7Ctri+Xtri+ nH2,A ) CmonoXmono + CdiXdi + Ctri+Xtri+

(16)

where C is the concentration of the respective compound class in the feed (mol/m3) and X is its corresponding total conversion (Xmono ) 0.2, Xdi ) 0.9, Xtri+ ) 1). The value obtained for the LGO feed used in this study was nH2,A ) 5. Under typical hydrotreating conditions (T ) 603 K, p ) 4 MPa, LHSV ) 2, HTO ) 300, LGO feed), the total aromatics conversion is around 50% (vide supra). We found under closely matching conditions (T ) 573 K, p ) 5 MPa, LHSV ) 1.9, HTO ) 510, LGO feed) a total sulfur conversion of 40%, as shown in Table 4. Because both HDS and HDA can be catalyzed by the same active sites on the catalyst, their conversion levels will generally be strongly coupled. Because of this, and because we have no other means to make a better estimation, we assume that the HDA conversion is equal to the HDS conversion. It is difficult to reliably estimate the molecular weight of the different aromatic compounds, as they may contain saturated side groups. As they are hydrogenated, their molecular weight will hardly change. We assume that all aromatic compounds have a similar molecular weight, equal to the average molecular weight of the feedstock (in Table 2). For their effective diffusivity, the value for naphthalene is used (Table 3). For the HDS reaction we make the following assumptions based on the literature:23,24 The major class of sulfur components in LGO is that of (substituted) (di)benzothiophenes. For simplicity, we assume that the molecular weight of all sulfur compounds is equal to that of dibenzothiophene. For their effective diffusivity, the value for anthracene is used (see Table 3). At intermediate sulfur conversion (below 80%) the ratedetermining reaction is the direct desulfurization (DDS) of dibenzothiophene analogues, which requires two H2 molecules to yield a biphenyl structure. Therefore, it is assumed that the reaction of one sulfur-containing molecule consumes two hydrogen molecules, i.e., the total hydrogen consumption per sulfur molecule converted nH2,S ) 2. In general, power law kinetics are used to describe the hydrodesulfurization (HDS) and aromatics hydrogenation (HDA) of realistic feedstocks. It should be noted that the apparent reaction order is usually larger than 1, as the feed consists of many classes of compounds with different reactivity. In this study we used a trickle bed microreactor, which can be described by a plug flow model. Hence, the kinetic rate constant for a first-order reaction was calculated using the following:

kV ) -LHSV

F0 ln(1 - X) F

(17)

where kV is the apparent or observed volumetric rate constant (1/h), LHSV is the liquid hourly space velocity (m3oil/m3cat/h), and F0/F is the correction factor for the difference in oil density at room temperature and at reaction temperature. For a reaction order other than 1 the following generalized equation was used:

kV ) LHSV

F0 1 [1 - (1 - X)1-n]Cin1-n F 1-n

(18)

in which n is the reaction order, Cin is the inlet concentration (mol/m3), and kV has the dimension (mol/m3)(n-1)‚h-1. The observed reaction rate per unit catalyst volume was calculated using the following:

rV,p )

Vbed Fp r ) Vbed kVCbn Vp V W

(19)

where Vbed is the catalyst bed volume and Vp is the catalyst volume (m3), which is a function of catalyst weight W (kg) and particle density Fp (kg/m3). Based on the properties of the feedstock used in this study, we can assume that the only significant reactions taking place are hydrodesulfurization (HDS) and aromatics hydrogenation (HDA). Both reactions consume hydrogen, which may cause hydrogen transport limitations. The hydrogen conversion and rate of hydrogen consumption were calculated according to the following:

Vm,H2,0 F HTO F0

XH2 ) (nH2,SCS,inXS + nH2,ACA,inXA)

rV,H2 ) nH2,SrV,S + nH2,ArV,A

(20) (21)

in which XS, XA, and XH2 are the conversions of sulfur, aromatics, and hydrogen, nH2,S and nH2,A are the average number of hydrogen molecules consumed per converted sulfur-containing or aromatic molecule, CS,in and CA,in are the inlet concentrations of sulfur and aromatics (mol/m3), Vm,H2,0 is the molar volume of hydrogen at standard conditions (0.024 m3/mol), HTO is the hydrogen to oil ratio (m3H2/m3oil), and rV,S, rV,A, and rV,H2 are the reaction rates of sulfur, aromatics, and hydrogen (mol/ m3s). We further assume that the liquid is always completely saturated with hydrogen. This implies that the gas to liquid transfer of hydrogen is much faster than the rate of hydrogen consumption and that the concentration of hydrogen is constant in the interparticle space. So, only external and internal mass transfer limitations are considered but not gas-liquid mass transfer. This was verified experimentally (not shown) by measuring the hydrogenation rate of naphthalene in our microflow system. The observed rate was 1 order of magnitude higher than during LGO processing under similar conditions and corresponded well to the naphthalene hydrogenation kinetics. Ronze et al.26 measured the mole fraction of hydrogen dissolved in straight run gas oil at complete saturation. From their work we obtain:

(

CH2 ) 0.134p exp -

953 F T M

)

(22)

where p and T are pressure (MPa) and temperature (K), CH2 is the concentration of hydrogen (mol/m3), F is the density (kg/ m3), and M is the molar mass (kg/mol) of the gas oil. The steadystate conversion levels, calculated using the above assumptions, are shown in Table 4. The catalyst effectiveness factors were obtained using these values. For all conditions, the calculated external effectiveness factors for the sulfur compounds and aromatics were larger than 95% for first-order kinetics and for second-order kinetics, as shown in Table 6. This means that no external gradients are present for the sulfur and aromatic compounds. The internal effectiveness factor is also larger than 95% for both sulfur and aromatics

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Table 6. External and Internal Catalyst Effectiveness Factors for First- and Second-Order Kinetics reactant

ηe (first order)

ηe (second order)

ηi (first order)

ηi (second order)

sulfur aromatics hydrogena

0.99-1.00 0.99-1.00 0.89-0.99

0.95-1.00 0.96-1.00 0.53-0.98

0.97-0.99 0.97-0.99 0.74-0.98

0.87-0.99 0.90-0.99 0.24-0.96

a The effectiveness factors for hydrogen are strongly dependent on the reaction conditions (see also Figure 9).

Figure 9. Internal effectiveness factor for hydrogen at the reactor inlet as a function of temperature and pressure for first- and second-order reaction (LGO, HTO > 500). Lines were drawn to guide the eye.

based on first-order kinetics, but in some cases it drops to 87% for sulfur and 90% for aromatics when second-order kinetics are assumed. Even for these lower values, there are no significant concentration gradients inside the catalyst pellets, as can be seen in Figure 1, and this will not lead to inhomogeneous coke deposition. Li, Chen, and Tsai27,28 showed that the effectiveness factor for HDS using a catalyst with similar structural properties as the one used in our study could be as low as 66% using heavy residue oil as feed. Lee et al.29 obtained an internal effectiveness factor of about 80% for the model reaction of indole using a similar catalyst as ours. The reason why effectiveness factors for the sulfur compounds and aromatics are above 90% in this study is that the effective diffusivities of these compounds in LGO are higher than in heavy residue oil and that the reaction rate is lower than in the model compound case. For hydrogen the case is different, as the concentration in the liquid is relatively low and a function of the hydrogen pressure in the gas phase. At the reactor inlet where the rate of hydrogen consumption is highest, the external effectiveness factor is still above 89% for all tests when we assume firstorder kinetics (see Table 6). For second-order kinetics, at low pressure (1 MPa) it drops to about 53%. The internal effectiveness factors are even lower and are plotted as function of

reaction conditions in Figure 9. It should be noted that the HDA reaction consumes a large portion of the available hydrogen. As both the HDA conversion and the hydrogen consumption were estimated based on the above assumptions, significant errors could arise in these values. The homogeneous coke distributions observed in the pellets aged at high pressure, and the nonuniform profiles observed at low pressure, confirm that diffusion limitations exist at low hydrogen availability. The typical bell shape of the profiles must therefore be related to the hydrogen concentration in the pellets. When we assume that the coke deposition rate is inversely proportional to the hydrogen concentration in the liquid (i.e., hydrogen suppresses coke formation) we can simulate the coke profiles by taking the reciprocal values of c from eq 8 and obtain the plots in Figure 10. Thus, under the assumption mentioned above, the coke profile is a projection of the hydrogen consumption rate inside the pellet, which is mainly a function of the aromatics hydrogenation rate (HDA). Especially the firstorder kinetic model results in the typical bell-shaped coke profile also observed in the real samples at low pressure, whereas for the second-order model the shape is slightly different. Nevertheless, when the internal effectiveness factors at 623 K from Figure 10 are used for the first- and second-order simulation (74% and 24%, respectively), the second-order simulation matches the real profile closest. Hence, if all of the above assumptions are indeed valid, the overall hydrogen consumption rate in the pellet can therefore best be approximated by a second-order dependency in hydrogen concentration. By using the internal effectiveness factor from the diffusivity equations (a value of 24% was found based on second-order kinetics), a very good match between the simulated and observed coke profile was found (compare Figures 8 and 10B). This indicates a satisfactory accuracy of the estimations made to obtain the diffusional properties and hydrogen consumption rates. As both could not be determined experimentally, the error in the calculated effectiveness factors may be large. However, the aim of these calculations was to show which reactants are possibly hindered by mass transfer limitations under hydrotreating conditions. By combining the observed coke profiles and calculated effectiveness factors, we can conclude that at low pressure, significant hydrogen diffusion limitations exist inside the catalyst pellet. It should be noted that these limitations are most severe at the inlet of the reactor. The catalyst pellets with the most extreme coke profiles probably came from the top of the bed. This can also explain why a lower effectiveness factor at the inlet has no measurable effect on the total conversion, as the latter is governed by the average effectiveness of the entire catalyst bed. When the hydrogen to oil ratio is lowered, the catalyst suffers from hydrogen depletion, which results in a

Figure 10. Simulated coke deposition profiles for different internal effectiveness factors (indicated above each curve) for first-order (A) and second-order reaction (B).

Ind. Eng. Chem. Res., Vol. 46, No. 2, 2007 429

lower conversion and more coke deposition. Summarizing, although the diffusion coefficient of hydrogen far exceeds those for the other reactants, only for hydrogen intraparticle gradients are significant during LGO processing. The reason lies in combination of the relatively low solubility of hydrogen and the positive overall order in hydrogen concentration. Conclusion During the artificial aging of the model NiMo/Al2O3 catalyst, no initial deactivation was observed. The amount of deposited coke was almost independent of pressure and temperature in the experiments at high hydrogen to oil ratio. At low HTO ratio and pressure the amount of deposited coke doubled. Under these conditions the catalyst suffers from hydrogen depletion, which results in a lower conversion and enhances coke deposition. Addition of polyaromatic compounds further decreased the conversion; however, it had no influence on the amount of coke. Apparently, the nature of the selected compounds in combination with their low concentration did not result in an enhanced coking tendency. The external and internal effectiveness factors for the sulfur-containing and aromatic reactants were close to or above 90% in all experiments, assuming first- or second-order reaction kinetics. The internal effectiveness factor for hydrogen is strongly dependent on the (hydrogen) pressure. The results show that at low pressure significant hydrogen concentration gradients exist inside the catalyst pellets due to mass transfer limitations. This leads to the formation of typical bell-shaped coke profiles, which could be explained by a simple model, assuming that the rate of coke deposition is inversely proportional to the hydrogen concentration inside the pellet. Up to about 25% of the initial surface area and about 30% of the initial pore volume was lost due to coke deposition. This had no significant effect on activity because the active sites were probably still accessible. A thin layer of coke deposited on the pore walls, and a small part of the pore structure was blocked by coke, of which a significant part deposits during the first hours on stream. The density of the coke was similar in all experiments, and poisoning of the active sites did not take place. Acknowledgment Albemarle Catalysts and NWO are gratefully acknowledged for their financial support. Literature Cited (1) Tanaka, Y.; Shimada, H.; Matsubayashi, N.; Nishijima, A.; Nomura, M. Accelerated deactivation of hydrotreating catalysts: Comparison to longterm deactivation in a commercial plant. Catal. Today 1998, 45, 319. (2) Marafi, M.; Stanislaus, A. Effect of initial coking on hydrotreating catalyst functionalities and properties. Appl. Catal., A 1997, 159, 259. (3) Yoshimura, Y.; Shimada, H.; Kubuta, M.; Nishijima, A.; Sato, T. Initial catalyst deactivation in the hydrotreatment of coal liquid over NiMo and CoMo-γAl2O3 catalysts. Appl. Catal. 1987, 21, 125. (4) Richardson, S. M.; Nagaishi, H.; Gray, M. R. Initial coke deposition on a NiMo/γ-Al2O3 bitumen hydroprocessing catalyst. Ind. Eng. Chem. Res. 1996, 35, 3940. (5) Begon, V.; Warrington, S. B.; Megaritis, A.; Charsley, E. L.; Kandiyoti, R. Composition of carbonaceous deposits and catalyst deactivation in the early stages of the hydrocracking of a coal extract. Fuel 1999, 78, 681. (6) Gualda, G.; Kasztelan, S. Coke versus metal deactivation of residue hydrodemetallization catalysts. Stud. Surf. Sci. Catal. 1994, 88, 145.

(7) Zeuthen, P.; Bartholdy, J.; Wivel, P.; Cooper, B. H. Formation of coke on hydrotreating catalysts and its effect on activity. Stud. Surf. Sci. Catal. 1994, 88, 199. (8) Weissman, J. G.; Edwards, J. C. Characterization and aging of hydrotreating catalysts exposed to industrial processing conditions. Appl. Catal., A 1996, 142, 289. (9) Dı´ez, F.; Gates, B. C.; Miller, J. T.; Sajkowski, D. J.; Kukes, S. G. Deactivation of a Ni-Mo/γ-Al2O3 catalystsInfluence of coke on the hydroprocessing activity. Ind. Eng. Chem. Res. 1990, 29, 1999. (10) Gray, M. R.; Zhao, Y.; McKnight, C. M.; Komar, D. A.; Carruthers, J. D. Coking of hydroprocessing catalyst by residue fractions of bitumen. Energy Fuels 1999, 13, 1037. (11) Hauser, A.; Stanislaus, A.; Marafi, A.; Al-Adwani, A. Initial coke deposition on hydrotreating catalysts. Part II. Structure elucidation of initial coke on hydrodemetallation catalysts. Fuel 2005, 84, 259. (12) Satterfield, C. N. Mass Transfer in Heterogeneous Catalysis; MIT Press: Cambridge, 1970. (13) Guisnet, M.; Magnoux, P. Organic chemistry of coke formation. Appl. Catal., A 2001, 212, 83. (14) Tsai, C. H.; Massoth, F. E.; Lee, S. Y.; Seader, J. D. Effects of solvent and solute configuration on restrictive diffusion in hydrotreating catalysts. Ind. Eng. Chem. Res. 1991, 30, 22. (15) Boned, C.; Ze´berg-Mikkelsen, C. K.; Baylaucq, A.; Dauge´, P. Highpressure dynamic viscosity and density of two synthetic hydrocarbon mixtures representative of some heavy petroleum distillation cuts. Fluid Phase Equilib. 2003, 212, 143. (16) Shanshool, J.; Hashim, E. T. Kinematic viscosity-temperature correlation for undefined petroleum fraction of a wide boiling range. Pet. Sci. Technol. 2001, 19, 257. (17) Furimsky, E.; Massoth, F. E. Deactivation of hydroprocessing catalysts. Catal. Today 1999, 52, 381. (18) Vogelaar, B. M.; Kagami, N.; van Langeveld, A. D.; Eijsbouts, S.; Moulijn, J. A. Active sites and activity in HDS catalysis: The effect of H2 and H2S partial pressure. Prepr. Pap.sAm. Chem. Soc., DiV. Fuel Chem. 2003, 48, 548. (19) Koltai, T.; Macaud, M.; Guevara, A.; Schulz, E.; Lemaire, M.; Bacaud, R.; Vrinat, M. Comparative inhibiting effect of polycondensed aromatics and nitrogen compounds on the hydrodesulfurization of alkyldibenzothiophenes. Appl. Catal., A 2002, 231, 253. (20) Wood, J.; Gladden, L. F. Effect of coke deposition upon pore structure and self-diffusion in deactivated industrial hydroprocessing catalysts. Appl. Catal., A 2003, 249, 241. (21) van Doorn, J.; Moulijn, J. A. A model of coke on hydrotreating catalysts under reaction conditions. Fuel Process. Technol. 1993, 35, 275. (22) Korre, S. C.; Klein, M. T.; Quann, R. J. Polynuclear aromatic hydrocarbons hydrogenation 1sExperimental reaction pathways and kinetics. Ind. Eng. Chem. Res. 1995, 34, 101. (23) Girgis, M. J.; Gates, B. C. Reactivities, reaction networks and kinetics in high-pressure catalytic hydroprocessing. Ind. Eng. Chem. Res. 1991, 30, 2021. (24) Girgis, M. J.; Gates, B. C. Catalytic hydroprocessing of simulated heavy coal liquids 1. Reactivities of aromatic hydrocarbons and sulfur and oxygen heterocyclic compounds. Ind. Eng. Chem. Res. 1994, 33, 1098. (25) Florusse, L. J.; Peters, C. J.; Pamies, J. C.; Vega, L. F.; Meijer, H. Solubility of hydrogen in heavy n-alkanes: Experiments and SAFT modeling. AIChE J. 2004, 49, 3260. (26) Ronze, D.; Fongarland, P.; Pitault, I.; Forissier, M. Hydrogen solubility in straight run gasoil. Chem. Eng. Sci. 2002, 57, 547. (27) Li, C.; Chen, Y. W.; Tsai, M. C. Highly restrictive diffusion under hydrotreating reactions of heavy residue oils. Ind. Eng. Chem. Res. 1995, 34, 898. (28) Tsai, M. C.; Chen, Y. W.; Li, C. Restrictive diffusion under hydrotreating reactions of heavy residue oils in a trickle-bed reactor. Ind. Eng. Chem. Res. 1993, 32, 1603. (29) Lee, S. Y.; Seader, J. D.; Tsai, C. H.; Massoth, F. E. Restrictive diffusion under catalytic hydroprocessing conditions. Ind. Eng. Chem. Res. 1991, 30, 29.

ReceiVed for reView July 6, 2006 ReVised manuscript receiVed September 14, 2006 Accepted November 20, 2006 IE060870A