Fuel 100 (2012) 17–23

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Early stage deactivation of heavy crude oil hydroprocessing catalysts

S.K. Maity a,⇑, E. Blanco b, J. Ancheyta a, F. Alonso a, H. Fukuyama c

a Instituto Mexicano del Petróleo, Eje Central Lázaro Cárdenas Norte 152, Col. San Bartolo, Atepehuacan, México, DF 07730, Mexicob École National Supérieur d’Ingénierie de Caen, 6 bld, Maréchal Juin, 14050 Caen, Francec Toyo Engineering Corporation, 8-1 Akanehama 2-Chome, Narashino-Shi, Chiba 275-0024, Japan

a r t i c l e i n f o a b s t r a c t

Article history:Received 8 June 2011Received in revised form 14 September 2011Accepted 5 November 2011Available online 23 November 2011

Keywords:HDMHDSDeactivationThiopheneHeavy oil

0016-2361/$ - see front matter � 2011 Elsevier Ltd. Adoi:10.1016/j.fuel.2011.11.017

⇑ Corresponding author.E-mail address: skumar@imp.mx (S.K. Maity).

Four different CoMo catalysts supported by alumina, alumina–titania, alumina–silica and carbon wereused to study the early stage of deactivation. Hydrodemetallization (HDM) and hydrodesulfurization(HDS) activities of these catalysts have been tested in high pressure, high temperature micro-plant byusing heavy crude oil as feed. Thiophene activity of the spent catalysts was compared with the freshand regenerated catalysts. The results show that the alumina and alumina–titania supported CoMo cat-alysts exhibit high and stable performance for both HDM and HDS activities. Though the initial activitiesof alumina–silica supported CoMo catalyst are high, the activities decrease rapidly with time-on-stream.Having higher acidic sites may cause this rapid deactivation. Thermogravimetric analysis results also sup-port that the nature of deposited carbon on this catalyst is different from the coke deposited on the otherthree catalysts. SEM-EDX results show that vanadium sulfides are more preferably deposited at the outersurface of carbon catalyst. Thiophene HDS activity of fresh, spent and regenerated results suggest that thedeactivation of alumina catalyst is cause of coke deposition whereas both metal sulfides and coke depo-sitions are responsible for CoMo/Al2O3–TiO2 and CoMo/Al2O3–SiO2 catalysts deactivation, particularly atthe early stage of hydroprocessing of the heavy oil.

� 2011 Elsevier Ltd. All rights reserved.

1. Introduction

Hydroprocessing catalyst for heavy crude oils is deactivatedrapidly during the reaction. The degree of deactivation dependson several factors like nature of crude oil, operating conditions,reactor design, etc. The deactivation is faster when the crude oilhas high percentage of asphaltenes and at severe reaction condi-tions. Hydroprocessing catalyst is mostly deactivated by coke andmetals deposits. To maintain the desired activity, the catalyst deac-tivation is compensated by continuously raising of the tempera-ture. After certain time, the deactivated catalyst needs to bereplaced by new one. Depending of the nature of crude oils andrequirement, the life of hydrotreating catalyst for heavy oil variesfrom 6 months to 1 year [1].

It has been proved that coke is formed very rapidly at earlystage of hydrotreating reaction and after that it becomes slow orconstant with time-on-stream. The coke is defined as soft and harddepending on its nature [2]. The ‘soft’ coke, which is formed at theinitial stage of reaction, is the principal cause of the loss of microp-ores and consequently surface area. It approximately reduces onethird of the porosity [3], whereas the ‘hard’ coke, which is formedin later stage of reaction, and metal sulfides take up the remainingporosity. At initial stage, coke is mostly deposited on the bare

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surface area. Its interaction with surface is much stronger whenthe feed contains more aromatic hydrocarbons. The soft coke hashigh hydrogen-to-carbon ratio whereas the hard coke contentslower H/C ratio. The coke formation is more when one feed con-tains high percentage of aromatics and heterocyclic compoundsthan other, though both having similar boiling ranges. Coke canbe produced from the pyrolysis of carbon or by precipitation ofasphaltenes molecules from heavy crude oil or/and as a result ofcondensation reaction of asphaltenes. Whatever origin it is, thecoke is deposited on pore mouth when catalyst pore is small,whereas it penetrates into the pore cavity when catalyst porediameter is comparatively large [4].

At early stage of reaction, the deactivation by metals is veryminimal. However with time this deposit increases and it becomescrucial in deactivation. In this respect vanadium and nickel are ofparticular concern because of the presence of these metals in highpercentage in heavy crude oils. The metals are commonlydistributed between porphyrin and nonporphyrin type of struc-tures [5–7]. These metal containing compounds are deposited intothe catalyst during hydrotreating and cause deactivation. Thedeactivation by metals is irreversible. Vanadium, due to its higherreaction rate, concentrates at the surface of the catalyst, whereasnickel is distributed more evenly throughout the catalyst pellet[8–10]. Vanadium can decorate the edge of a molybdenumdisulfide slab just as can a nickel or cobalt promoter. However,vanadium displaces nickel from the edge sites on the molybdenum

18 S.K. Maity et al. / Fuel 100 (2012) 17–23

slab and is then a source of deactivation since promotion by vana-dium is less than that by nickel [11].

The reason and mechanism of catalyst deactivation have beeninvestigated by various researchers. Some of them assumed thatrapid coke deposit might cause initial deactivation of the catalysts,while others reported that deposition of metals should also be con-sidered for initial deactivation. It was also stated that deactivationmight be the composite effect of metals and coke deposits and itwas not easy to distinguish quantitatively between deactivationby coke and by metals. Though there are several investigationson the reasons of deactivation, the role of support on this processis not studied thoroughly. Therefore, in this work, early stage ofdeactivation is investigated on different supported catalysts. Alu-mina, alumina–titania, alumina–silica and carbon supported CoMocatalysts have been tested by Maya crude HDT and thiophene HDS.

2. Experimental methods

2.1. Preparation of catalysts

The catalysts were prepared by incipient wetness technique andco-impregnation method. The appropriate amount of ammoniumheptamolybdate and cobalt nitrate salts were dissolved into dis-tilled water and made a clear solution. The amount of water takenfor making salt solution was just sufficient to fill up the pore ofsupport. This clear solution was impregnated on dry support andimpregnated samples were allowed overnight at room tempera-ture. Samples were then dried 7 h at 120 �C and calcined at450 �C for 5 h in presence of air. For the carbon supported catalyst,the calcination was performed at 450 �C in presence of nitrogen.

2.2. Characterization of catalyst

BET specific surface area, pore volume and pore size distributionof fresh, spent and regenerated catalysts were measured by nitro-gen adsorption at �196 �C (Automatic Micromeritics ASAP 2100).The percentage of carbon was also measured on the spent cata-lysts. The spent catalysts were washed with hot toluene by Soxhletprocess and dried at 110 �C before carbon and metal analysis. Cokeis defined in this work as being carbon content on a spent catalyst.The metals distribution of catalysts was measured by a scanningelectron microscope (SEM), model XL30ESEM, Philips. The spentcatalysts were also characterized by thermogravimetric analysis(TGA). Weight loss by combustion was measured by Perkin-Elmermodel TGA 7-HT. In this experiment, around 50–100 mg of catalystwas heated from room temperature to 1000 �C at a rate of 10 �C/min in the presence of air (flow 50 mL/min).

2.3. Catalyst activity test

Heavy oil hydrotreating activity tests were performed in a highpressure fixed-bed micro-reactor in up flow mode. The oxide cata-lysts were sulfided in situ before actual run was started. Ten milli-liters of oxide catalyst was loaded with equal volume of diluent,carborandum (0.2 mm size). Both catalyst and carborandum weremixed and divided into five parts. Each part of the mixture wasloaded into the reactor at a time and tapped little bit. The catalystwas then dried for 2 h at atmospheric pressure at 120 �C. After dry-ing, the catalyst was allowed to soak for 2 h at 150 �C. Light gas oil(LGO) was used for soaking. This light gas oil contains 1.7 wt% ofsulfur. Actual sulfiding agent was introduced after soaking. Thesulfiding agent was light gas oil with dimethyl disulfide (DMDS,1 wt%). Sulfidation was done at 28 kg/cm2 pressure at two differenttemperatures. The first sulfidation was done at 260 �C for 3 h andfinally the catalyst was sulfided at 320 �C for 5 h. The experimental

conditions are: total pressure, 54 kg/cm2; reaction temperature,380 �C; liquid hourly space velocity (LHSV), 1.0 h�1; hydrogen-to-hydrocarbon ratio, 356 m3/m3. A mixture of Maya heavy oil withhydrodesulfurized naphtha (50/50 wt/wt) was used for catalystactivity tests. The first balance was taken after stabilization of9 h. The products were collected at the duration of 12 h each.

Thiophene hydrodesulfurization activity of fresh, spent andregenerated catalysts was tested in micro-plant. In this plant,0.5 g of fresh catalyst was sulfided in situ into a glass-tubular reac-tor at atmospheric pressure and 400 �C temperature. For sulfida-tion, hydrogen was passed through a container having CS2. Thesaturated mixture of CS2 and hydrogen was passed through thereactor. Hydrogen flow was 50 mL/h and duration of sulfidationwas 2 h. After sulfidation, the catalyst was flushed at this temper-ature by H2 until no CS2 could be detected in the effluent gas. Thi-ophene feed was introduced through the gas bubblers. H2 flow ratewas 50 mL/min. Reaction products were analyzed by an on-line gaschromatography using a FID detector. The spent catalyst afterobtaining from high pressure micro-plant was washed and it wastested for HDS of thiophene. Before actual run, the spent catalystwas also sulfided as describe above. HDS activity of the spent cat-alysts was tested at period of 4 h. This same spent catalyst wasregenerated by control flow of air at 500 �C and sulfided beforetesting its activity.

The total amount of metals in the feed and products weremeasured by Atomic Absorption (Thermoelectron model SolaarAA). Sulfur was analyzed by X-ray fluorescence (HORIBA modelSLFA-2100).

3. Results and discussion

3.1. Physicochemical properties of feed and catalysts

In this work we used a mixture of Maya heavy crude with hyd-rodesulfurized naphtha (1:1 wt) as a hydrotreating feed for micro-plant. Sulfur, nitrogen, asphaltene, nickel, and vanadium contentsof this feed are 1.86, 0.253, 7.96 wt%, 26.17 and 123.35 wppm,respectively. All catalysts have 10 wt% of MoO3 and 3 wt% of CoO.Four different supports (alumina, alumina–titania, alumina–silicaand carbon) have been used to prepare CoMo catalysts. Both titaniaand silica were 10 wt% on the alumina support. The physical prop-erties of these catalysts are given in Table 1. The carbon supportedCoMo catalyst has very high surface area of 1288 m2/g. Total porevolume of this catalyst is also high, 1.15 mL/g. Alumina supportedcatalyst has the highest average pore diameter. The specific surfacearea, pore volume and pore diameter of alumina–titania andalumina–silica supported catalysts are in intermediate betweenalumina and carbon supported catalysts.

3.2. HDM and HDS activities with real feed

The hydrodemetallization activities of different supported cata-lysts are presented with time-on stream (TOS) in Fig. 1. All HDMactivities decrease with TOS; however the trend is different for dif-ferent catalysts. Alumina–titania supported CoMo catalyst showsthe highest HDM activity. It is also observed that though the initialactivity of alumina–silica catalyst is high, the activity decreasesrapidly with TOS. Alumina supported catalyst has also high HDMactivity and it is stable with TOS. Carbon supported CoMo catalystshows the lowest activity. The HDM activity of four supportedCoMo catalyst is in the order of: CoMo/Al2O3–TiO2 > CoMo/Al2O3 > CoMo/Al2O3–SiO2 > CoMo/C.

The HDS activity has also been studied for heavy crude and theresults are given in Fig 2. Similar to HDM, CoMo/Al2O3–TiO2 cata-lyst has the highest HDS activity and CoMo/C shows the lowest.

Table 1Physical properties and coke and vanadium depositions on catalysts.

Properties CoMo/Al2O3 (A) CoMo/Al2O3–TiO2 (B) CoMo/Al2O3–SiO2 (C) CoMo/C (D)

F S R F S R F S R F S

SSA (m2/g) 309 248 308 312 270 300 416 242 404 1288 39TPV (mL/g) 0.71 0.37 0.69 0.46 0.37 0.48 0.72 0.26 0.67 1.15 0.12APD (Å) 92 0.60 91 60 55 64 69 43 66 36 135PSD (V%)<50 Å 28.47 72.49 29.5 49.66 71.17 53.12 38.8 70.81 58.44 59.04 12.0850–100 Å 31.99 14.73 44.99 38.82 17.51 34.86 53.32 14.84 34.24 12.68 9.85100–200 Å 30.31 8.69 20.66 6.78 6.75 7.73 4.62 7.03 4.95 13.55 37.5200–500 Å 5.38 4.09 3.77 3.17 3.04 3.14 2.23 4.37 1.84 13.86 35.39500–1000 Å 0.85 0 1.09 1.57 1.53 1.15 1.03 2.95 0.53 0.87 5.18>1000 Å 3.00 0 0 0 0 0 0 0 0Coke (wt%) 13.2 6.1 10.4 -Vanadium (wt%) 1.11 0.15 0.12 0.09

F = fresh, S = spent, R = regenerated.

Fig. 1. Heavy oil HDM activity of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B), CoMo/Al2O3–SiO2 (C) and CoMo/C (D) catalysts.

S.K. Maity et al. / Fuel 100 (2012) 17–23 19

In this case also rapid fall of HDS activity is noticed for alumina–silica supported catalyst.

The rate of deactivation for all catalysts is also calculated byusing following equation [12,13].

Xt ¼ X0e�btn

where Xt is the conversion at time t, X0 is the initial conversion, b isthe deactivation rate constant, n = 1 [13].

Fig. 2. Heavy oil HDS activity of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B), CoMo/Al2O3–SiO2 (C) and CoMo/C (D) catalysts.

�ln(Xt/X0) is plotted against TOS for HDM and HDS reactions inFigs. 3 and 4, respectively. Although the total reaction time is low(60 h); even within this time different deactivation rate is noticed.The deactivation rates (b value) have also been calculated from thestraight lines. The HDM deactivation rates are 0.0047, 0.0076,0.0222 and 0.0078 for the catalyst A, B, C and D, respectively whilethese rates are 0.0036, 0.0041, 0.0357, 0.0076 for HDS reaction. It isfound from the above values that the deactivation rate for HDMreaction is marginally faster than that of HDS reaction. The metalcompounds in the heavy crude oils are very large complex mole-cules. On the other hand sulfur containing molecules in heavy frac-tions are in general two types-one is bigger in size, attached withasphaltene structure, and the other is smaller in size, non-asphaltenic. During hydroprocessing reaction the solubility of theasphaltenes is reduced by removing their aliphatic chains, particu-larly at initial period when the catalyst activity is high causingasphaltenes to precipitate from the product and hence they aredeposited on the catalysts. As a result, catalyst is deactivated bythis deposition. Therefore, the deactivation rate for HDM reactionis faster than that for HDS because metals are associated withasphaltene moiety. The smaller size of non-asphaltenic sulfur com-pounds is still having path to enter into the pore cavity and hencethe deactivation rate is slower [14,15]. It is also noted that the rateof deactivation of the CoMo/Al2O3–SiO2 catalyst is very high forboth HDM and HDS reactions. It suggests that the alumina–silicasupport may have some acidic sites and these sites enhance theformation of coke on the bare support surface. The deposited cokeon the spent catalyst given in Table 1 is high, around 10 wt%. The

Fig. 3. Rate of HDM deactivation of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B), CoMo/Al2O3–SiO2 (C) and CoMo/C (D) catalysts.

Fig. 4. Rate of HDS deactivation of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B), CoMo/Al2O3–SiO2 (C) and CoMo/C (D) catalysts.

20 S.K. Maity et al. / Fuel 100 (2012) 17–23

lowest deactivation rates for both HDM and HDS reaction arefound for carbon supported catalysts. The carbon is neutral supportand hence it shows low activity and therefore, the coke depositionis also low (Table 1). This may be the reason for showing stableperformance with TOS.

3.3. Thiophene HDS activities on fresh, spent and regenerated catalysts

The hydrodesulfurization activities of fresh, spent and regener-ated catalysts were studied and the results are given in Fig. 5. It isworth to mention that the spent catalysts are obtained from highpressure micro-plant where a mixture of Maya crude with naphthawas treated for 60 h. The HDS activity of regenerated catalyst D isnot presented here, because carbon supported catalyst cannot beregenerated by oxidation. Fig. 5 shows that the catalyst A has veryhigh HDS activity. Though HDM and HDS activities of heavy oil arelow for carbon supported catalyst D, thiophene HDS activity ismoderately high. At least it is higher than those of catalysts Band C. HDS activity of catalyst A is drastically reduced due to thedeactivation. However, when coke in the spent catalyst is burnedoff, the catalyst literally gains its original activity. The other twocatalysts B and C show almost similar trend, i.e. the fresh catalysthas higher HDS activity than spent and regenerated catalysts. Eventhough catalyst D has moderate activity, its spent catalyst showsvery low activity. Thiophene HDS activity of the fresh sulfidedCoMo catalysts is in the order of: CoMo/Al2O3 > CoMo/C > CoMo/Al2O3–TiO2 � CoMo/Al2O3–SiO2. The overall results indicate thatthe different supported catalysts have different HDS activity. Thedeactivation trend is also different for different catalysts. The cokedeposition may have adverse effect on alumina and carbon sup-ported catalyst. However, the metals sulfide may also take part

Fig. 5. Thiophene HDS activity of fresh (F), spent (S) and regenerated (R) catalysts.

at early stage of deactivation for the other two supported catalysts(B & C).

3.4. Scanning electron microscopy (SEM)

The profiles of vanadium and coke deposits are also studied bySEM–EDX and their radial distributions on the spent catalysts arepresented in Figs. 6 and 7, respectively. Fig. 6 shows that vanadiumis evenly deposited throughout the catalyst particle for the\catalysts A, B and C. These catalysts are supported by Al2O3,Al2O3–TiO2 and Al2O3–SiO2, respectively. However, more vanadiumdeposition occurs at outer surface of the carbon supported catalyst.The coke deposition profiles for the catalysts A, B and C are more orless homogeneous throughout the catalyst particle (Fig. 7). Sincethe catalyst D is supported by carbon, the carbon deposited duringthe hydroprocessing reaction cannot be distinguished from carbonsupport and hence the coke deposition prolife is not presented forcatalyst D in this figure.

3.5. Themogravimetric analysis (TGA)

Catalyst weight loss due to combustion of the spent catalystshas been measured by thermogravimetric analysis. The weight lossand its derivative with temperature are presented in Fig. 8A–D. Allderivative curves clearly show two principal weight losses ataround 100 �C and at 400 �C. The second peak at around 400 �C isvery sharp and prominent. The weight loss at first peak is for theloss of water and the second peak is identified as the weight lossdue to coke burning. However, the second peak position is notthe same for all catalysts. The peak position for the catalysts Aand B is very near to 400 �C, whereas this peak has been shiftedto higher temperature for the catalysts C and D. It is 460 and440 �C for the catalysts C and D respectively. The catalyst D is car-bon supported catalyst and carbon (or coke) has also been depos-ited during hydroprocessing reaction. So the net weight loss ofthis catalyst during TGA analysis is due to the carbon from supportand from coke. To distinguish these two sources of carbon, TGAanalysis is also performed on fresh CoMo/C catalyst and the resultsare compared with spent catalyst in Fig. 8D. The nature of thederivative curves for spent catalyst D is very wide compared withthe other catalysts and with its fresh catalyst. It indicates that bothpeaks of support carbon and coke from reaction are superimposedone over other. The combustion of fresh CoMo/C catalyst is startedat 290 �C and ended at 550 �C whereas the combustion of spentCoMo/C catalyst is started at 260� and ended at 575 �C. A very wide

Fig. 6. Radial distribution of vanadium of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B),CoMo/Al2O3–SiO2 (C) and CoMo/C (D) spent catalysts.

Fig. 7. Radial distribution of coke of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B), CoMo/Al2O3–SiO2 (C) and CoMo/C (D) spent catalysts.

S.K. Maity et al. / Fuel 100 (2012) 17–23 21

peak is also observed at higher temperature for the catalysts A, Band C. These wide broad peaks appear around 900 �C, 650 �C and750 �C respectively. This broad peak is absent for carbon supportedcatalyst D. The characterization of coke by TGA experiment hasalso been studied by various researchers [16,17]. Begon et al.[16] observed three different regions during combustion of spenthydrotreating catalyst. The region I is <180 �C, region II is in

Fig. 8. TGA curves of CoMo/Al2O3 (A), CoMo/Al2O3–TiO2 (B)

between 180 and 330 �C and region III is in between 330 and750 �C. The fist region is obviously due to loss of water, whereasthe region II is owed to the weight loss of mobile carboneous res-idues which is defined as soft coke. The bulky hard coke is com-busted at higher temperature and that appears at region III. TGAexperiment on the spent catalysts has been studied in detail bySahoo et al. [17]. They have also used mass spectroscopy to analyzethe gaseous products coming out due to burning of spent catalysts.The authors have found that peak at 350–360 �C was due to theweight loss of presulfiding material present in the catalyst andbulk carboneous materials is burned off at 480 �C. They have alsoobserved other peaks due to leaving out of SO2 from the samples.In this study, we observed a sharp peak at 400–460 �C and whichis the principal peak for combustion of carboneous material. Thenature of this coke is not similar for all the catalysts. For the cata-lysts A and B the coke is combusted at lower temperature (400 �C)whereas it is high (460 �C) for the catalyst C. So it can be stated thatfor alumina–silica supported catalyst, the deposited carbon is rela-tively hard in nature. On the other hand, for carbon supported cat-alyst the presence of both soft and hard coke is found from Fig. 8D.

3.6. Physical properties of spent catalysts

The specific surface area (SSA), total pore volume (TPV), averagepore diameter (APD) and pore size distribution (PSD) of fresh,spent and regenerated catalysts are compared in Table 1. The poressize distribution of each supported catalyst has also been compared

, CoMo/Al2O3–SiO2 (C) and CoMo/C (D) spent catalysts.

22 S.K. Maity et al. / Fuel 100 (2012) 17–23

in Fig. 9A–D. Table 1 shows that the specific surface area and totalpore volume of the spent catalysts are reduced with respect to cor-responding fresh catalysts. The loss of these physical properties isnot same for all the catalysts. The surface area of the spent catalystD has dramatically reduced from 1288 to 39 m2/g within a periodof 60 h reaction with heavy crude oil. A least change of the surfacearea is observed for the catalyst B. The catalyst A regains its surfacearea and total pore volume by regeneration of the spent catalyst.When we compare the HDS activity of fresh and regenerated cata-lyst A, it is also noticed that both catalysts show almost similaractivity.

The pore size distribution of the spent catalyst has been chan-ged compared with fresh catalyst. The pore volume of the spentcatalysts below 50 Å increases, whereas pore volume having higherpore diameter decreases. When the changes of total pore volume ofspent catalysts are compared it is noticed that total pore volume of

Fig. 9. Pore size distribution of fresh (F) and spent (S) and regenerated (R) CoMocatalysts.

catalyst D is drastically reduced from 1.15 to 0.12 mL/g, whereas itis the least for the catalyst B. Fig. 9B shows that alumina–titaniasupported catalyst has strong ability to resist the changes of itstextural properties. The surface area, total pore volume and its poredistribution of this catalyst do not change significantly duringhydroprocessing reaction of heavy crude, particularly at early stageof reaction.

Alumina–silica supported catalyst has more coke deposition ascompared with alumina–titania supported catalyst. The formermay have more acidic sites that enhance cracking of the apshaltenemolecules and hence coke formation. Our micro-plant activity re-sults also support that though the initial activities of alumina–silica supported CoMo catalyst is high; they rapidly decrease withTOS. Fig. 9C also reveals that the pore size distribution of catalyst Chas been also changed considerably upon deactivation. The alu-mina supported catalyst A shows very high HDS and HDM activi-ties. The coke and vanadium depositions on this catalyst are alsovery high. Our thiophene HDS activity results indicate that thespent catalyst substantially loss activity. However, when this cata-lyst is regenerated, the catalyst recuperates almost all HDS activitylike a fresh catalyst. So it suggests that the coke deposit mainlycauses deactivation of this catalyst. The pore size distributions offresh, spent and regenerated catalysts show that the pore structureof catalyst A changes dramatically with deactivation; however, thepore structure becomes almost similar to the fresh one when thecarbon was burned off from the spent catalyst (Fig. 9A). Table 1shows that both coke and vanadium depositions are high on thespent catalyst A. Thus it suggests that the change of pore structureis mainly due to the coke formation. Due to the coke deposition,the activity of this catalyst is also drastically reduced and whenthis coke is eliminated by burning, the catalyst recuperates its ori-ginal activity. It also hints that though metal deposition is high, itmay not take part on deactivation.

The deactivation for other two supported catalysts (Al2O3–TiO2

and Al2O3–SiO2) is moderate. Thiophene HDS activity of the regen-erated catalysts (B and C) is lower than the corresponding freshcatalyst. It reveals that coke and metal deposits take part for deac-tivation and reduce HDS activity for the catalysts B and C. It sug-gests that the early stage of deactivation is unavoidable and thecauses of it may vary from catalyst to catalyst.

The catalyst at the beginning of a reaction is very active andhence conversion is high. At the early stage of reaction, coke is pre-dominantly deposited on the external surface of the catalyst. Indue course of operation, coke gradually penetrates into the interiorcavity of the catalyst particle. Therefore, deactivation can occur byboth phenomena-pore mouth plugging and core poisoning-depending on the catalyst pore structure. For a small pore, catalystdeactivation occurs via pore mouth plugging. On the other hand,for a large pore, the catalyst is deactivated by the core poisoning.However, the metals from the feed are gradually deposited onthe catalyst surface. At the early stage of hydrotreating reactionwhen the metals deposition is very low, metals sulfides may coverthe interior active sites of the particle rather than pore mouthplugging and it causes catalyst deactivation [18,19]. So, from ourresults we noticed that both coke and metals depositions areresponsible for early stage of deactivation. However, the influenceof each may not be the same for different supported catalysts. Theadverse effect of metals deposit is absent for alumina supportedcatalyst. Whereas both coke and metals take part on the deactiva-tation of alumina–titania and alumina–silica supported CoMo cat-alysts. The reason of catalyst deactivation for carbon supportedcatalyst is not clear, however dramatic change of pore structureis observed for this catalyst and it is mainly due to the coke depo-sition. It is also noted from SEM analysis that vanadium sulfidesmay deposit on the surface due to its smaller pore diameter.Gualda and Katsztelan [20] observed that both metal sulfides and

S.K. Maity et al. / Fuel 100 (2012) 17–23 23

coke deactivated HDM performance of NiMo/Al2O3 catalyst butdeactivation by metals is more prominent even at the initial stageof operation.

4. Conclusion

The causes for early stage deactivation are studied on differentsupported catalyst by using heavy crude oil. Four different sup-ports (alumina, alumina–titania, alumina–silica and carbon) areused to prepare CoMo catalysts. Thiophene HDS activity of freshCoMo/Al2O3 catalyst is the highest and it is reduced dramaticallyfor the spent catalyst. However, the regenerated catalyst shows al-most equal activity as fresh catalyst does. The pore size distribu-tions of fresh and regenerated catalysts also show that bothcatalysts have almost similar pore distribution. It suggests thatfor this catalyst the cause of catalyst deactivation is mainly cokedeposit. On the other hand, both coke and metals sulfides deposi-tions cause the deactivation of CoMo/Al2O3–TiO2 and CoMo/Al2O3–SiO2 catalysts. The HDM and HDS activities of carbon supportedcatalyst are the lowest among all the catalysts studied in this work.However, its thiophene HDS performance is better than the othertwo mixed oxide supported catalysts and it is only lower than alu-mina supported catalyst. Moreover, the spent carbon supportedcatalyst does not show any thiophene activity.

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