New insights in catalyst structure, activity and stability
Deactivation of Hydroprocessing Catalysts
Deactivation of Hydroprocessing Catalysts
New insights in catalyst structure, activity and stability
Proefschrift
ter verkrijging van de graad van doctor aan de Technische Universiteit Delft,
op gezag van de Rector Magnificus prof. dr. ir. J.T. Fokkema, voorzitter van het College voor Promoties,
in het openbaar te verdedigen op maandag 21 februari 2005 om 13:00 uur door Bastiaan Maarten VOGELAAR
scheikundig ingenieur geboren te Leidschendam
Prof. dr. J.A. Moulijn
Samenstelling promotiecommissie:
Rector Magnificus, voorzitter
Prof. dr. J.A. Moulijn, Technische Universiteit Delft, promotor Prof. dr. F. Kapteijn, Technische Universiteit Delft
Prof. dr. J.A.R. van Veen, Technische Universiteit Eindhoven
Prof. dr. R. Prins, Eidgenössische Technische Hochschule Zürich Prof. dr. E. Payen, Université des Sciences et Technologies de Lille Dr. A.D. van Langeveld, Technische Universiteit Delft
Dr. S. Eijsbouts, Albemarle Catalysts, Amsterdam
Prof. ir. J. Grievink Technische Universiteit Delft (reservelid)
This research was financially supported by Albemarle Catalysts and the Netherlands Organisation for Scientific Research (NWO).
ISBN: 90 6464 730 5
Printed by Ponsen & Looijen B.V., Wageningen Copyright © 2005, B.M. Vogelaar
Chapter 1: Introduction 1 Chapter 2: Coke profiles in commercial spent hydroprocessing catalysts 17 Chapter 3: Simulation of coke and metal deposition in catalyst pellets 33
using a non-steady state fixed bed reactor model
Chapter 4: Emergence of coke profiles and diffusion limitations during 63 artificial aging of hydroprocessing catalysts
Chapter 5: Catalyst deactivation during thiophene HDS: 83 The role of structural sulfur
Chapter 6: Identification of edge structures and thiophene HDS active 103 sites on Mo and NiMo sulfide catalysts
Chapter 7: New insights in the reaction mechanism and active sites 129 involved in dibenzothiophene HDS over a NiMo catalyst
Chapter 8: Evaluation 149 Samenvatting 159 Publications 161 Presentations at conferences 163 Dankwoord 165 Curriculum Vitae 167
Introduction
Hydroprocessing is one of the key unit operations in today’s oil refineries. Its importance will increase in the coming decades, due to more stringent fuel quality legislations and further growth of the global fuel demand. In this chapter, the basics of hydroprocessing, and the need for further research and development, especially in the field of hydroprocessing catalysts, will be briefly discussed. The research goals and approach, for the work presented in this thesis, will be defined.
The modern refinery
Oil refineries produce a broad range of transportation fuels and petrochemical raw materials, which are obtained by processing and blending the appropriate oil fractions. Figure 1 shows a simplified process scheme of a typical oil refinery [1]. Crude oil is fed to an atmospheric distillation column to yield several fractions of different boiling range: LPG and gas, straight run gasoline, naphtha, middle distillate or straight run gasoil (SRGO), and heavy atmospheric gas oil (HAGO). The bottom product, atmospheric residue, is fed to a vacuum distillation column to obtain vacuum gas oil (VGO) and vacuum residue. Gasoline and diesel have by far the largest demand, hence, refiners want to maximize the gasoline and diesel yields.
Ideally, gasoline has a high hydrogen to carbon atomic ratio (H/C = 2.25 if we assume isooctane as reference), no heteroatoms (except maybe some oxygen), a low volality and a high octane number [2]. Highly branched paraffins with 8-10 carbon atoms would best fulfill all these requirements. In contrast, the ideal diesel fuel would consist of linear paraffins, with a carbon atom range from 10 to 20. A limited number of branched chains would be beneficial in winter to prevent freezing. In diesel, (poly-) aromatic compounds and heteroatoms should be eliminated.
A substantial fraction of the crude oil (HAGO, VGO and vacuum residue) is unsuitable for blending into the gasoline and diesel product streams directly. These high boiling fractions are too heavy in terms of molecular weight, have a too low degree of saturation (H/C ratio) and contain relatively large amounts of heteroatoms and (poly-) aromatic species. Several catalytic processes
reduce the molecular weight, “hydrogen in” or “carbon out” to increase the H/C ratio, isomerization and alkylation to make branched molecules, and heteroatom removal.
Figure 1 shows the location of the three main conversion units in the refinery: the reformer, the fluid catalytic cracker (FCC) and the hydrocracker (HC). For simplicity, this scheme does not show processes to convert vacuum residue, e.g. coking, visbreaking etc. The reforming process is used to increase the aromatic and olefinic content of straight run gasoline, resulting in a higher octane number [3]. A high-value by-product of this process is hydrogen, which has a crucial role in the modern refinery (see below). The role of the reforming process will gradually shift towards petrochemicals production when future legislations will limit the aromatic (benzene) content of gasoline. Alternative processes, such as isomerization and alkylation, have been developed to produce isoparaffins with a high octane number. However, in contrast to reforming, these processes do not produce hydrogen. Fluid catalytic cracking (FCC) is the key conversion unit for many modern refineries [4]. Heavy molecules are “cracked” into lighter ones using an acid catalyst (zeolite) at high temperature. A fraction of the oil feed is converted into coke deposits, which are combusted to CO2 in the regenerator of the FCC unit and generates the heat required for the
cracking process. Hence, FCC can be considered a “carbon out” process. Due to the nature of the cracking reactions the product stream is rich in aromatics, olefins and branched molecules. Therefore, the main product of FCC is gasoline. FCC also produces light olefins that can be sold as raw materials to the polymer industry. Hydrocracking (HC) is a cracking process mainly developed for the production of diesel. Because the acid cracking catalyst is combined with a hydrogenation catalyst, the product contains mainly paraffinic molecules in the diesel boiling range. Hydrocracking is thus a typical “hydrogen in” process, and requires a substantial amount of (expensive) hydrogen.
Hydrodesulfurization or HDS is a unit operation found in many locations in the refinery process scheme (Figure 1). Its main function is the catalytic removal of sulfur from the various oil streams, but also other heteroatoms, like nitrogen, oxygen and metals, can be removed using similar catalysts. These processes, hydrodesulfurization (HDS), hydrodenitrification (HDN), hydrodeoxygenation (HDO), hydrodemetallization (HDM), and also aromatics hydrogenation (hydrodearomatization or HDA), all require hydrogen and share the common name hydroprocessing or hydrotreating. Their original purpose is to protect down-stream acid (FCC, HC) and metal (reforming) catalysts against poisoning by sulfur, nitrogen and metal compounds (HDS units 2 and 4 in Figure 1). The reformer (see above) can supply the hydrogen needed for these processes. Nowadays, the maximum sulfur content specifications of gasoline and diesel are the main reason for refiners to perform deep hydrodesulfurization (see below). Therefore, virtually every process stream needs to be hydroprocessed before entering the blending section (HDS units 1, 3, 5 and 6 in Figure 1). These processes require more hydrogen than can be supplied by the reformer, and forces modern refineries to implement hydrogen plants, which produce hydrogen by steam reforming of heavy residues or natural gas [5].
Table 1: Some typical heteroatom and aromatic compounds found in petroleum fractions [6]. Compound class Examples
R-SH R-S-R’ R-S-S-R’
Sulfur compounds
Thiols (mercaptans), sulfides and disulfides
Thiophenes, benzothiophenes, and dibenzothiophenes
Nitrogen compounds
Pyrrole, indoles and carbazoles
Pyridine, quinolines and acridines
Oxygen compounds
Furan, carboxylic acids and phenols
Aromatics
Benzene, tetralin and biphenyl
Naphthalene and anthracene
Phenanthrenes and pyrene
Metal compounds Porphyrins
Hydroprocessing technology
Since the introduction of the first commercial hydroprocessing reactors in 1925, this technology has been the subject of extensive research. Topsøe et al. [6] give a thorough review of the developments in this field up to 1996. Table 1 lists the typical classes of compounds found in oil fractions, which are to be converted by hydroprocessing. The abundance of these molecules in the different fractions depends on their boiling points: Untreated gasoline mainly contains thiols, sulfides, disulfides, thiophenes, pyrroles, pyridines, furans, phenols and phenyls. The other compounds in Table 1 are mainly found in the heavier fractions like straight run gas oil (diesel), HAGO and VGO. These compounds are more difficult to hydroprocess than the lighter compounds present in untreated gasoline. For example, the hydrodesulfurization rate of sulfur compounds over the same catalyst decreases in the order: thiols and (di-) sulfides > thiophenes > benzothiophenes > dibenzothiophene (DBT) > 4,6-dimethyl dibenzothiophene [2]. Especially 4,6-substituted DBTs are very difficult to desulfurize, presumably because of the shielding of the sulfur atom by the two substituents.
Girgis and Gates [7] performed extensive research on the reactions occurring in hydroprocessing. The generally accepted reaction mechanism for the HDS of dibenzothiophenes is shown in Figure 2. Two different steps are involved: hydrogenation and C-S bond breaking (hydrogenolysis). The order of these steps determines the reaction pathway: In the hydrogenation (HYD) route the aromatic ring is first (partially) hydrogenated, followed by sulfur removal. In the direct desulfurization route (DDS) the C-S bonds are broken without initial hydrogenation of the molecule. In particular, the DDS route is more difficult for 4,6-substituted DBTs, probably due to steric hindrance. When one of the aromatic rings is hydrogenated, the molecule is deformed, making it easier for the catalyst to attack the C-S bonds.
Figure 2: Proposed reaction network for the HDS of dibenzothiophene; (a) hydrogenation steps, (b) C-S bond breaking steps [6].
Hydroprocessing is commonly performed in a high-pressure trickle bed reactor [8]. Catalyst pellets are stacked in a packed bed, and gas (hydrogen) and liquid (oil) are fed co-currently from the top (Figure 3). The reactor operates in the so-called trickle flow regime, in which the catalyst pellets are fully wetted with the liquid, and both gas and liquid flow along the pellet external surface (insert Figure 3). Gas and liquid are separated at the bottom of the reactor for further processing. Generally, the HDS reactor off gas is split into a hydrogen-rich stream that is recycled, and an H2
S-rich stream that is sent to a sulfur recovery unit (Claus plant) [3]. HDS catalysts come in various shapes (Figure 3), but cylindrical and multi-lobe extrudates are most common [9]. Typical pellet diameters are 1 to 3 mm, with a length of several millimeters. Multi-lobe extrudates are preferred over cylinders or spheres for their lower pressure drop and better gas-liquid contacting and mass transfer characteristics. However, the drawback of using complex-shaped pellets is the reduced catalyst loading per unit reactor volume, due to the larger void space between the pellets [6].
A typical hydroprocessing reactor is several meters in height and diameter, with a catalyst volume of 30 to 100 m3. Typical operating conditions for different hydroprocessing feeds are summarized in Table 2 [6]. Clearly, the hydroprocessing of heavier feeds requires more severe process conditions (temperature, pressure) and longer contact times (lower space velocity). Because heavier feeds contain more aromatic species and heteroatoms, more hydrogen is needed for their processing. The total hydrogen feed rate is usually at least five times higher than the consumption rate, to maintain a high H2 partial pressure in the lower section of the catalyst bed, and to suppress
the build-up of gaseous inhibiters like H2S and ammonia. gas in liquid in gas out liquid out trickle flow gas in liquid in gas out liquid out trickle flow
sphere pellet ring
cylinder trilobe quadrulobe
sphere pellet ring
cylinder trilobe quadrulobe
Figure 3: Schematics of a trickle bed reactor, adapted from [8], and catalyst particle shapes used for HDS [9].
Hydroprocessing catalysts
The major part of today’s hydroprocessing units uses conventional CoMo or NiMo based catalysts, whose active metal composition has hardly changed since their first industrial application in 1925 [9]. It was found that highly dispersed molybdenum sulfide exhibits a low, but stable catalytic activity in a sulfur-containing environment. By adding the so-called promoter elements cobalt or nickel, in an atomic Co/Mo or Ni/Mo ratio of 0.3 to 0.5, a ten-fold increase in activity could be achieved [6]. Further improvements were made by the addition of phosphorus, which is claimed to modify the support interaction, and by optimization of the preparation method.
The catalysts are prepared by wet impregnation of a support material, usually γ-Al2O3 in the
form of extrudates. The impregnation solution contains Mo, Co and/or Ni and other additives. After drying and calcination the oxidic precursor is obtained. This precursor is activated in-situ in the hydroprocessing reactor by sulfidation. In most cases the oil feed is spiked with a sulfur-containing agent like dimethyldisulfide (DMDS) as sulfiding medium, or the catalyst is impregnated with a sulfur-containing agent (e.g. organic or inorganic polysulfides). The sulfiding medium is converted into H2S at elevated temperature and H2 pressure, which transforms the oxidic precursor into the
active sulfidic catalyst.
Many researchers investigated the structure of the active catalyst [6]. Using a broad range of (quasi) in-situ analysis techniques, a number of solid phases could be identified on the catalyst under working conditions (Figure 4): Small clusters of MoS2 and Co9S8 or NiS, also called bulk
sulfides, Co and Ni atoms that have migrated into the support (forming aluminates), and a unique phase, called CoMoS or NiMoS. This phase contains Co or Ni, Mo and S and has a morphology similar to MoS2. By correlating the activity of unpromoted MoS2 to its cluster size, it was
concluded that the active sites for HDS are located on the edges of the MoS2 crystallites. These
crystallites consist of hexagonal layers of Mo atoms, sandwiched between two layers of sulfur atoms. These sulfur atoms shield the Mo atoms from interaction with the reactants, but at the edges
Table 2: Typical process conditions and hydrogen consumption for various hydroprocessing operations, adapted from [6].
Feed Temperature (K) H2 pressure (MPa) LHSV1 (h-1) H(Nm2 consumption 3/m3 oil) Naphtha 593 1 – 2 3 – 8 2 – 10 Kerosine 603 2 – 3 2 – 5 5 – 15 SRGO 613 2.5 – 4 1.5 – 4 20 – 40 VGO 633 5 – 9 1 – 2 50 – 80 Atm. residue 643 – 683 8 – 13 0.2 – 0.5 100 – 175 VGO HC2 653 – 683 9 – 14 1 – 2 150 – 300 Residue HC2 673 – 713 10 – 15 0.2 – 0.5 150 – 300 1) Liquid hourly space velocity (m3
oil/m3cat/h) 2) Hydrocracking
of the crystal, sulfur atoms may be removed, creating so-called vacant sites. It is generally believed that these vacant sites interact with the heteroatoms of the reactants, followed by hydrogenolysis.
Kinetic modeling studies give evidence for two different catalytic sites, one for hydrogenation (HYD route) and one for C-S bond breaking (DDS route), shown in Figure 2. Daage and Chianelli [10] proposed that the HYD sites are located only on the top and bottom “rim” of unsupported stacked MoS2 clusters, whereas the DDS sites are on the entire edge. The physical structure of these
sites, the actual catalytic mechanism and the role of the promoter (Co or Ni) is however still unclear.
The role of the promoter in CoMo and NiMo based catalysts is widely discussed. Two main theories have been developed; the CoMoS model proposed by Topsøe and co-workers [11], and the contact synergy model suggested by Delmon [12]. Delmon and co-workers envisage the CoMoS phase as a close interaction of small cobalt sulfide particles and MoS2, in which the promoter atoms
activate and spill over hydrogen to the vacant sites on the MoS2 edges, which perform the
hydrogenolysis reactions. The Topsøe group considers the CoMoS phase as a unique new structure, in which the promoter atoms decorate the MoS2 edges and form highly active sites. Together with
the University of Aarhus, the Topsøe group has recently given new indications for the validity of their CoMoS theory, using new high resolution scanning tunneling microscopy (STM) techniques on a gold-supported model catalyst [13]. Based on the STM results, they claim that the cobalt atoms are built into the MoS2 edge structure (Figure 5). Recent advances in the field of computational
simulations have enabled researchers to show the thermodynamic feasibility of such structures [14]. However, many details about the nature of the active sites and the catalytic mechanism have not yet been elucidated. It might even be true that both the CoMoS phase and the bulk sulfides (contact synergy) play a role in industrial hydroprocessing catalysts [15].
Figure 4: Schematic view of the different phases present in a typical alumina-supported HDS catalyst, adapted from [6].
Future trends and the need for research
In the last decades, SOx emissions from transportation fuels have been identified as an important
contributor to acid rain, and legislative limitations have been imposed on the sulfur content of gasoline and diesel. The policy makers (governments and international councils) base their decisions on the current state of the art in hydroprocessing and the expected technological advances in this field. The United States, Japan and Europe are the first to introduce ultra-low sulfur specifications for gasoline and diesel (Table 3). Diesel fuel sulfur levels should be reduced to below 10-15 ppm within the next four years [16, 17]. Other countries are expected to follow these regulations as soon as the technology to reach these low sulfur limits has evolved further.
Figure 5: STM image of a real CoMoS nanocluster on Au(111), and proposed ball model of the structure, both taken from [13] (top view, light spheres = S, dark spheres = Mo, grey spheres = Co).
Table 3: Planned maximum sulfur specifications for gasoline and diesel fuel in the US, Japan and Europe, adapted from [16, 17].
Year Gasoline specification (ppm S) Diesel specification (ppm S)
US Japan Europe US Japan Europe
Current 200 200 150 500 500 350 2004 50 2005 50 2006 30 15 2007 30 10 2008 10 10
Because gasoline is much easier to desulfurize than diesel (see above) and the global gasoline demand is growing at a slower rate than the diesel demand (Figure 6), refineries will be able to meet the ultra-low sulfur targets for gasoline without major investments or increased operational costs. Only in Northern America gasoline has a larger market share than diesel. In Europe, the total gasoline demand is decreasing, while the diesel demand grows due to the increased popularity of diesel-fueled cars [18]. The market share for fuel oils is decreasing worldwide, forcing refiners to upgrade heavy fractions for blending into the diesel pool. Unfortunately, these lower-quality feedstocks will require severe hydroprocessing to meet the future diesel specifications. The refining industry is therefore facing a double-edged challenge: to meet more stringent specifications for diesel products, while producing more diesel using feedstocks of lower quality [19].
Refineries may choose from several technological solutions to reach the ultra-low sulfur limits for diesel. Table 4 compares the different options that are currently available for industrial application [20]. Typically, a 1.5 times higher activity is needed to lower the product sulfur level from 2500 ppm to 500 ppm, without shortening the catalyst lifetime (i.e. without increasing the reaction temperature). Further reduction to 50 or 10 ppm requires a 4 to 5 times higher activity [19]. Changing the process conditions alone (temperature, H2 pressure, purity or feed rate) is clearly not
enough to reach the 10-15 ppm target while maintaining high throughput and diesel yield (Table 4). Higher activity catalysts offer the best solution in terms of effectiveness, ease of implementation and costs. As a result, the hydroprocessing catalyst market is projected to grow $34 million per year (4.3% annual growth) in the coming years [21].
0 10 20 30 40
Gasoline Diesel and
Kerosine
Fuel oil Lubes and
Waxes
M
illion barrels per da
y North AmericaAsia - Pacific
Europe Rest of world 0 10 20 30 40
Gasoline Diesel and
Kerosine
Fuel oil Lubes and
Waxes
M
illion barrels per da
y North AmericaAsia - Pacific
Europe Rest of world
Figure 6: World oil demand in 1997 per region and product type, adapted from [18]. Extension lines show total estimates for 2007 based on 1997 growth rates.
Hydroprocessing catalyst manufacturers work hard to develop new high-activity catalysts, which enable refineries to reach the new sulfur targets, without having to build additional reactors. Figure 7 shows the advances made by three leading catalyst manufacturers; Haldor Topsøe, Criterion and Albemarle Catalysts (formerly Akzo Nobel). The reported values are based on their publications [19, 22, 23], and normalized to the individual 1980-1989 benchmarks. No information was available to compare absolute numbers between different manufacturers. However, the relative catalyst activities achieved by each manufacturer are quite similar up to 2002. Before 1995, most commercial catalysts were of the standard or “Type I” kind. The discovery and commercialization of the “Type II” catalyst led to an activity doubling between 1995 and 2002. Type II catalysts are prepared using a chelating complex, such as NTA, to improve the dispersion and promoter effectiveness [24]. Criterion’s Centinel and Albemarle’s Stars catalyst families are based on this technology. In 2002, Akzo Nobel Catalysts (now Albemarle) commercialized their proprietary Nebula catalyst, which has an over four times higher activity than the 1980-1989 benchmark. It is the first commercial catalyst able to produce ultra-low sulfur diesel in a hydroprocessing unit designed for 500 ppm sulfur production. The catalyst is not made by the traditional impregnation of alumina carriers, but uses a new support material developed by ExxonMobil and Akzo Nobel [23]. The actual composition is not disclosed, but also these new catalysts contain Co, Ni and Mo as main active components.
Today’s industrial solutions for ultra-deep hydrodesulfurization rely on conventional molybdenum sulfide based catalysts. A comprehensive understanding of the catalytic mechanisms and nature of the active sites is still lacking, in spite of the many insights obtained through R&D
Table 4: Comparison of technological options to meet 10-15 ppm sulfur diesel specifications [20].
Option Effective-ness Ease of implemen-tation Cost indication Comment
Use improved catalyst + ++ ++ Depends on catalyst vendors
Adjust feed cut point + ++ ++ Loss of diesel yield
Raise reactor temperature - ++ ++ Shortens catalyst life,
increases unit downtime Improve reactor
efficiency
+/- + - Improvement depends on
existing reactor design Add more catalyst
volume
+/- +/- +/- Install additional reactor or replace the existing one
Increase H2:oil ratio +/- +/- +/- Requires additional recycle
compressors Remove H2S in recycle
gas
+/- +/- - Requires new amine
scrubber or hot stripping Increase H2 partial
pressure
- +/- - Requires H2 purification or imported H2
Install two-stage system ++ - - Revamp existing 1-stage to
2-stage; add new reactor
Install new grassroots unit ++ -- -- Most expensive but most
effective
over the past years. Many advances in hydroprocessing catalyst development are the result of optimizations and incidental discoveries. Alternative desulfurization technologies (adsorption, extraction, oxidation, alternative catalysts, biocatalysis, counter-current reactors) are being developed, but most of them will not be commercially attractive or available on the short term [1, 25, 26]. Therefore, the elucidation of the mechanism of conventional sulfide-based catalysts can still be a highly rewarding challenge in hydroprocessing research, and drive the short-term development of highly active and stable CoMo and NiMo catalysts.
Catalyst deactivation
Not only the activity but also the stability of the catalyst is extremely important in hydroprocessing, as it determines how long the reactor can be operated before the catalyst needs to be regenerated or replaced. Future trends in sulfur target levels and feedstock quality may force refineries to operate their hydroprocessing units at more severe conditions (i.e. higher temperatures and heavier feeds), which may adversely affect the catalyst lifetime. Furimsky and Massoth [27] recently published an in-depth review on the deactivation of hydroprocessing catalysts.
Figure 8 shows a schematic view of the different processes occurring during catalyst deactivation. At the high temperatures required in hydroprocessing, the active metal clusters may agglomerate because of solid-state diffusion. This reduces the accessible active surface area for the reactants. For the same reason, the promoter may segregate from the CoMoS-type phase, to form less active bulk sulfides. Strongly adsorbing molecules (like nitrogen compounds), or carbonaceous
TK -5 5 0 TK -5 5 4 TK -5 5 4+ TK -5 74 TK -5 74 Cen ti n el Cen ti n el Std. Sta n da rd Ty pe 10 0 Nebula Stars CoMo CoMo Sta rs 0% 100% 200% 300% 400% 500% 1980-1989 1990-1994 1995-1999 2000-2001 2002-2004 R e la tive vo lu m e ac tivi ty ( R V A ) Haldor Topsøe Criterion Albemarle base TK -5 5 0 TK -5 5 4 TK -5 5 4+ TK -5 74 TK -5 74 Cen ti n el Cen ti n el Std. Sta n da rd Ty pe 10 0 Nebula Stars CoMo CoMo Sta rs 0% 100% 200% 300% 400% 500% 1980-1989 1990-1994 1995-1999 2000-2001 2002-2004 R e la tive vo lu m e ac tivi ty ( R V A ) Haldor Topsøe Criterion Albemarle base
Figure 7: Advances in HDS catalyst development by three leading manufacturers over the past 25 years, adapted from [19, 22, 23]. Note that the relative activities were normalized to the individual 1980-1989 values (base).
accessibility. The pore structure of the catalyst may be blocked by coke or metal deposits from the feed, preventing reactants from reaching the active sites inside the catalyst pellet. Coke and metal deposition depend strongly on the feedstock composition. Finally, the active sites may undergo structural changes not related to the above processes, reducing the catalytic activity. It is difficult to study the catalyst structure under reaction conditions, but a few researchers have given evidence that such changes may occur [28, 29]. Still, many details about this process are unknown.
Aim of the research
Hydroprocessing catalysts show a fast deactivation due to coke deposition and adsorption of poisons during the first few days on stream, after which a long period of almost stable activity is attained (Figure 9). The activity drops again fast due to blocking of the pore structure by metals and coke at the end of the run, which can be up to two years of operation. These processes accelerate during the run, because in practice, the reactor temperature is gradually increased to compensate for the loss in activity. The catalyst needs to be regenerated when the maximum temperature of operation is reached [27]. Especially the initial deactivation offers a challenge for catalyst improvement; up to 50% of the initial activity may be lost within a few days on stream. If answers
? (a) (b) (c) (d) (e)
Figure 8: Possible deactivation mechanisms for hydroprocessing catalysts: (a) sintering, (b) promoter segregation, (c) change of active phase, (d) poisoning, (e) pore blocking (white = promoter, grey =
MoS2, black = deposits, hatched = support).
Time on stream Act iv ity initial deactivation (days) steady state (0.5-2 years) end of run
can be found to the question why the catalyst deactivates initially, changing the catalyst properties or operating conditions may enhance its long-term performance significantly.
The aim of the research presented in this thesis is to obtain a better insight in the role of the different deactivation mechanisms in the initial deactivation of conventional hydroprocessing catalysts. The focus is on the effects on the catalytic performance and the structure of the active phase. The secondary aim of the research is to obtain more detailed insights in the HDS reaction mechanism and the structure of the active sites.
Figure 10 shows the approach followed to reach the above goals. The research is divided into five sub-topics: pore blocking, coke deposition, sintering and segregation, (changes of) the active phase, and the catalytic reaction mechanism. Pore blocking is investigated by analyzing the coke deposition profiles in industrial spent catalysts (Chapter 2), modeling of the observed deposition profiles (Chapter 3) and artificial aging of a model catalyst using an industrial feedstock (Chapter 4). Poisoning by coke deposition is studied by aging a model catalyst under industrial (Chapter 4) or model conditions (Chapter 5). In Chapter 5, also sintering and segregation of the promoter are studied. The structure and stability of the active phase is analyzed using a dedicated in-situ analysis technique under gas phase (Chapter 6) and liquid phase conditions (Chapter 7). Finally, the relations between the catalytic reactions and active phase are evaluated, as reported in Chapter 7.
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29. B.M. Vogelaar, P. Steiner, A.D. van Langeveld, S. Eijsbouts and J.A. Moulijn, Appl. Catal. A:
Coke profiles in industrial spent hydroprocessing catalysts
The distribution of coke inside the pellets of several industrial spent hydroprocessing catalysts was analyzed using Raman spectroscopy. The shape of the coke profiles gives us information on the deactivation mechanism and the magnitude of diffusion limitations. Samples taken at the end of run showed the most pronounced coke profiles. Typical M-shaped profiles observed in a guard bed catalyst suggest a sequential deactivation mechanism and a strong diffusion resistance. Consequently, it is concluded that near the end of life the catalyst effectiveness factor can be significantly reduced by coke deposition.
Introduction
An important aspect in catalytic reactor design is the catalyst effectiveness factor [1]. When a catalyst is applied in a trickle bed reactor, it requires a number of macroscopic properties like a high mechanical strength, a low pressure-drop and a good gas-liquid contacting efficiency (i.e. a high external surface area). To meet these requirements hydroprocessing catalysts are typically used as spheres, pellets, beads or extrudates. Most common are cylindrical extrudates (solid or hollow) with a diameter of one to three millimeters and several millimeters in length, often with a ribbed external surface (e.g. trilobes and quadrulobes, see Figure 1). As a consequence the characteristic diffusion length is relatively large, which may lead to concentration gradients inside the catalyst extrudates, thus lowering the catalyst effectiveness factor [2]. How much the catalyst effectiveness is actually reduced depends on the diffusional properties of the feedstock molecules. Heavy feedstocks like vacuum gas oil (VGO) will cause more problems than lighter feeds like light gas oil (LGO). Of course the diffusion rates can be enhanced by modification of the pore structure of the catalyst; larger pores will give a higher overall effectiveness.
Several studies have been performed on the diffusional characteristics of hydroprocessing catalysts, in particular on catalysts of which the pore structure was modified (degraded) by coke deposition [3-6]. Catalysts that have a high effectiveness factor at the start of the run may develop significant diffusion resistance at longer time on stream when part of the pore structure is blocked by deposited coke. The presence of diffusion limitations leads to the formation of non-uniform coke profiles in the catalyst pellet. In fact, the shape of the coke profile gives us information on the deactivation mechanism and the magnitude of diffusion limitations. If we could monitor the coke deposition process inside the catalyst pellet, we would be able to probe the condition of the catalyst’s pore structure and the extent of diffusion limitations within the catalyst extrudate. This concept was applied to a series of spent catalysts used in industrial refinery hydroprocessing units and laboratory test reactors under different conditions.
Many analysis techniques are suitable for the analysis of coke profiles in catalyst particles, like nuclear magnetic resonance imaging (NMR), scanning Auger microprobe analysis (SAM), nuclear microprobe analysis (NRA, ERA, PIXE and RBS), electron microprobe analysis (EPMA) and Figure 1: Cylindrical, trilobe and quadrulobe extrudate geometry.
this reason Raman spectroscopy was chosen as analytical tool for this study. Nowadays, Raman equipment is relatively cheap and very easy to operate. Of course this technique has it drawbacks, like problems with fluorescence, sensitivity or quantification. Fluorescence is a phenomenon that can completely obscure the Raman spectrum, and may be caused by organic impurities, basic surface OH groups, proton superpolarizability or reduced transition metal ions when resonantly excited [13]. All these conditions may be present in the case of spent hydroprocessing catalysts. Fortunately, current technology has greatly improved the sensitivity of Raman spectrometers by using highly sensitive CCD cameras and high performance holographic notch filters. However, quantitative Raman analysis is still a difficult task because the scattering cross-sections and the local morphology of many materials are not known, and may change as a function of analysis conditions like temperature and pressure. Moreover, Raman cross-sections of surface species, like ad-layers or adsorbates, may change due to interaction effects with the support. For heterogeneous, polydisperse materials, large errors may arise due to different particle sizes and morphologies [13]. Consequently, care must be taken when interpreting absolute Raman intensities as a measure for the total amount of coke. We can however assume that within the particle, the morphology is more or less the same. Hence, relative differences in the particle, i.e. coke profiles, can be measured with reasonable accuracy.
Raman spectroscopy is a widely used technique for studying carbons and coke samples. One of the first attempts to explain the Raman spectra of graphite-like materials was published in 1970 by Tuinstra and Koenig [14]. Graphite is a layered structure and, as the interaction between the stacked layers is very weak, Raman phonons are only active within a single layer. In a perfect single crystal of graphite only one Raman band is observed at 1575 cm-1, which is called the G band or graphitic band. This band is due to the so-called E2g stretching mode (Figure 2), which involves the in-plane
bond-stretching motion of pairs of sp2 carbons. This mode does not require the presence of six
membered rings but occurs at all sp2 sites, as it does in aromatic and olefinic molecules. In those cases the frequency always lies in the range of 1500-1630 cm-1 [15]. In amorphous graphite another major contribution in the Raman spectrum is observed around 1355 cm-1: the D or disordered band [14]. This is the A1g “breathing” mode of the aromatic ring clusters in the graphite sheets (Figure 2).
In an infinite plane of graphite (i.e. in a single crystal) this mode is forbidden, as localized breathing vibrations will be neutralized by vibrations with opposite amplitude in the same plane. Only
E2gmode A1gmode
breathing of the graphitic plane as a whole will yield a Raman signal, which requires this plane to be sufficiently small. The smaller the size of the graphitic planes, the larger the Raman contribution in the 1355 cm-1 region will be [14].
For purely graphitic cokes this interpretation appears to be straightforward, however, theoretical and experimental studies revealed the presence of a disorder-induced Raman band at 1620 cm-1 called D’ band, which in many cases interferes with the G band [16, 17]. Even more complexity is introduced when the coke is partially hydrogenated; studies on amorphous hydrogenated carbons show that the D and G modes are also sensitive to the type of bonding, i.e. sp2 or sp3 [18]. New and more powerful computational methods enable the simulation of Raman spectra of large polycyclic aromatic hydrocarbons, which can be envisaged as molecular subunits of graphite [19]. These results explain the appearance of the D band in disordered graphitic materials and, moreover, that the possible content of hydrogen is irrelevant for the intensity of the D band. In many cases the structure of the carbon species is not purely graphitic, like in activated carbons or coked catalysts.
Several researchers attempted to characterize these materials using a two-band [20-22], three-band [18], or four-three-band fitting [23] of the Raman spectra. When there is significant line broadening of the Raman signal, as is often the case in these highly amorphous samples, peak fitting becomes very error-sensitive, and interpretation will be difficult. Because of this, and because we expect the coke to be quite ill-defined, we did not attempt to deconvolute the coke contribution in the Raman spectra in this study. Only the total carbon intensities were calculated by integrating the Raman bands between 950 and 1750 cm-1, yielding the total area of the D and G bands. This procedure was used to obtain the coke distribution profiles in industrial spent catalyst pellets.
Laser source Spot size control CCD detector Sample Notch filter Microscope Grating Camera Laser source Spot size control CCD detector Sample Notch filter Microscope Grating Camera
Experimental
Sample preparation
The spent catalysts were provided by Akzo Nobel Catalysts, Amsterdam. The history of each sample is listed in Table 1. All samples were commercial NiMoP alumina supported catalysts with a 1.3 mm diameter quadrulobe shape, except for catalyst A1, which was 3 mm in diameter. The samples had been unloaded from the reactor and stored immersed in oil (except series A) to prevent degradation by air. The samples of series A had already been dried before storage. Prior to analysis the oil was removed by Soxlet extraction in oxygen-free toluene during six hours. The samples were subsequently dried in an air flow at room temperature to evaporate the toluene. This is a proven method to remove oil residue from spent hydroprocessing catalysts [24]. Akzo Nobel Catalysts also provided the oxidic precursors or “fresh catalysts” of sample series A, for comparison to their spent counterparts.
SEM/EDX analysis
Scanning Electron Microscopy (SEM)-analysis was conducted using a Philips SEM XL20 equipped with a SED detector for secondary and back-scattered electrons. Catalyst A1 and A2 were selected because they could be analyzed without pretreatment (toluene washing), as the pellets of these samples were dry. Several extrudates of each sample were cut into small pieces of about one millimeter thickness, mounted on an aluminum stub using adhesive carbon tape and sputtered with gold to provide sufficient electrical conductivity. Further details on magnification and settings are indicated in the micrographs. To measure the metal distribution in the fresh catalyst extrudates, several line scans were made using EDX (energy-dispersive analysis of X-rays).
Raman analysis
Raman analysis was performed using a Renishaw Ramascope System 2000 instrument linked to a Leica microscope (Figure 3). A 514 nm, 20 mW Ar+ laser was used as excitation source. The backscattered light was filtered for Rayleigh scattering using a holographic notch filter. The spectrograph uses a grating to disperse the light over the CCD detector, which records the Raman spectrum with a resolution of 4 cm-1. To protect the coked samples from oxidizing, the analyses were carried out under protective conditions, that is, in a Linkam THMS600 flow cell connected to a 100 ml/min stream of dry nitrogen. Samples were prepared by cutting several extrudates into small pieces of about one millimeter thickness, and sticking them onto a glass support disk using double-sided adhesive tape. A minimum of five pellets per sample was analyzed, which is recommended to obtain a representative average of the data [9]. The Raman mapping procedure was fully automated; sample positioning and laser focusing were handled by a Prior H101 motorized XYZ-stage connected to the instrument.
Tabl e 1: H istor y of spe nt c atal yst s use d i n thi s study. Ca ta ly st Time on s tr eam A pplicatio n Feed 1 Temperature (start of run) H2 Pre ssur e LHSV 2 HT O 3 Comments A1 1 week Pilot test VGO 633 K 7.0 M Pa 1 400 Standard s iz e (1.3 mm ) A2 1 week Pilot test VGO 633 K 7.0 M Pa 1 400 L ar ge extrudates (3 mm) B1 <3 months Industrial V GO 653 K 11 MPa 0.6 1000 Earl y s hutdown B2 <3 months Industrial VGO 653 K 11 MPa 0.6 1000 Emerg enc y shutdown C1 12 m onths Industrial H VGO 684 K 5.7 M Pa 0.7 380 St
andard pore siz
e C2 12 m onths Industrial H VGO 684 K 5.7 M Pa 0.7 380 W ide p ore catal ys t D1 7 m onths Guard bed C GO 573 K 8.0 M Pa 8.5 * >350
Standard pore siz
e D2 7 m onths Guard bed C GO 573 K 8.0 M Pa 8.5 * >350 W
ide pore catal
yst 1 ) VGO = v ac uum g as oi l, HVGO = he av y va cuum ga s oil, CGO = cok er ga s oil 2 ) L H SV = liquid hourl y sp ace v elocit y (m 3 oi l / m 3 cat / h ) 3 ) HTO = h ydro gen to oil fl ow ratio (standard m 3 H2 / m 3 oil )
* ) Based on the catal
ys t v ol ume of the gu
ard bed onl
y Tabl e 1: H istor y of spe nt c atal yst s use d i n thi s study. Ca ta ly st Time on s tr eam A pplicatio n Feed 1 Temperature (start of run) H2 Pre ssur e LHSV 2 HT O 3 Comments A1 1 week Pilot test VGO 633 K 7.0 M Pa 1 400 Standard s iz e (1.3 mm ) A2 1 week Pilot test VGO 633 K 7.0 M Pa 1 400 L ar ge extrudates (3 mm) B1 <3 months Industrial V GO 653 K 11 MPa 0.6 1000 Earl y s hutdown B2 <3 months Industrial VGO 653 K 11 MPa 0.6 1000 Emerg enc y shutdown C1 12 m onths Industrial H VGO 684 K 5.7 M Pa 0.7 380 St
andard pore siz
e C2 12 m onths Industrial H VGO 684 K 5.7 M Pa 0.7 380 W ide p ore catal ys t D1 7 m onths Guard bed C GO 573 K 8.0 M Pa 8.5 * >350
Standard pore siz
e D2 7 m onths Guard bed C GO 573 K Tabl e 1: H istor y of spe nt c atal yst s use d i n thi s study. Ca ta ly st Time on s tr eam A pplicatio n Feed 1 Temperature (start of run) H2 Pre ssur e LHSV 2 HT O 3 Comments A1 1 week Pilot test VGO 633 K 7.0 M Pa 1 400 Standard s iz e (1.3 mm ) A2 1 week Pilot test VGO 633 K 7.0 M Pa 1 400 L ar ge extrudates (3 mm) B1 <3 months Industrial V GO 653 K 11 MPa 0.6 1000 Earl y s hutdown B2 <3 months Industrial VGO 653 K 11 MPa 0.6 1000 Emerg enc y shutdown C1 12 m onths Industrial H VGO 684 K 5.7 M Pa 0.7 380 St
andard pore siz
e C2 12 m onths Industrial H VGO 684 K 5.7 M Pa 0.7 380 W ide p ore catal ys t D1 7 m onths Guard bed C GO 573 K 8.0 M Pa 8.5 * >350
Standard pore siz
e D2 7 m onths Guard bed C GO 573 K 8.0 M Pa 8.5 * >350 W
ide pore catal
yst 1 ) VGO = v ac uum g as oi l, HVGO = he av y va cuum ga s oil, CGO = cok er ga s oil 2 ) L H SV = liquid hourl y sp ace v elocit y (m 3 oi l / m 3 cat / h ) 3 ) HTO = h ydro gen to oil fl ow ratio (standard m 3 H2 / m 3 oil )
* ) Based on the catal
ys t v ol ume of the gu
ard bed onl
Results
SEM/EDX
Figure 4 shows a SEM micrograph of the catalyst precursor of sample A2. The structure of the extrudates is essentially homogeneous, although some cracks are visible which may have been caused by the cutting procedure. However at a greater magnification numerous spots can be observed, which apparently have a much more open structure than the rest of the catalyst material (Figure 4B).
The corresponding spent catalyst did not show any different features at low magnification (Figure 5A). When we look more up close (Figure 5B) the surface appears to have more cracks, which could be interpreted as some sort of dried wax-like or coke layer.
A B
Figure 4: SEM micrographs of the oxidic catalyst precursor (3 mm extrudates) and magnification of the inclusions (B).
A B
Figure 5: SEM micrographs of spent catalyst sample A2 (ex pilot plant) and magnification of the cracked structure (B).
Commercial spent catalyst A1 exhibited a quite different picture when observed under the SEM. On the surface of the cross-section, next to the inclusions described above, many volcano-like “bumps” can be seen, with a highly cracked surface (Figure 6).
Figure 7 shows a typical EDX line scan (spatial resolution 2 µm) of the nickel and molybdenum content, for the NiMoP fresh catalyst precursor. Apart from a few small deviations, both the Ni and Mo distribution are homogeneous throughout the particle.
Raman
Figure 8A shows a Raman spectrum of spent catalyst A1 taken without any sample pretreatment. This is a typical example of fluorescence; a huge “bump” that extends all over the spectral window and obscures most of the Raman bands. This phenomenon is caused by species in the sample that backscatter photons of random energy upon irradiation. One way to overcome this problem is to reduce the laser output power, or in this case by de-focusing the beam. Figure 8B
A B
Figure 6: SEM micrographs of spent catalyst sample A1 (ex pilot plant) and magnification of the typical protrusions (B). Ni(L)/Al(K) Mo(L)/Al(K) 0.0 0.2 0.4 0.6 0.8 1.0 -500 -250 0 250 500 r (microns) 0.0 0.2 0.4 0.6 0.8 1.0 -500 -250 0 250 500 r (microns) 0.0 0.2 0.4 0.6 0.8 1.0 -500 -250 0 250 500 r (microns) 0.0 0.2 0.4 0.6 0.8 1.0 -500 -250 0 250 500 r (microns) Figure 7: SEM/EDX profiles of Ni and Mo of the NiMoP catalyst precursor.
even more spectral features become visible. The Raman spectrum was compared to that of the freshly sulfided catalyst and its oxidic precursor in Figure 9, where the features are indicated in the graph by their corresponding wavenumbers. Unfortunately the de-focusing technique was not effective for all samples. The toluene washing procedure however proved effective to reduce fluorescence in all spent catalyst samples, indicating that fluorescence may be caused by oil residue in the catalyst pores.
0 1000 2000 3000 4000 Raman Shift (cm-1) 0 1000 2000 3000 4000 Raman Shift (cm-1) Intensity (a.u.) 0 1000 2000 3000 4000 Raman Shift (cm-1) 0 1000 2000 3000 4000 Raman Shift (cm-1) A B
Figure 8: Raman spectra of spent catalyst A1, focused (A) and de-focused (B).
100 300 500 700 900 1100 1300 1500 1700 1604 1340 408 385 350 220 916 954 578 408 385 Oxidic Sulfided Spent Raman shift (cm-1) R am an si gn al ( a.u .)
The inclusions that were observed in the SEM micrographs are also visible using the optical microscope, as shown in Figure 10A. These spots have a darker color than the rest of the catalyst cross-section, which might indicate a higher coke content. However, when the intensity of the coke bands was mapped using Raman (in air), the actual coke content of the spots appeared to be lower (Figure 10B). When we look at the mapped area again after the analysis (Figure 10C), small white spots have emerged at the positions where the surface was probed by the laser beam. Apparently, the surface of the sample has changed during the analysis, possibly by burning off the coke. This shows the high reactivity of the coke when Raman analysis is performed in air. Therefore, in subsequent experiments, the samples were protected under a nitrogen atmosphere.
Coke profiles
For a detailed view of the coke distribution, several areas of spent catalyst A1 were mapped with a high spatial resolution (10 µm). As an example Figure 11A shows a micrograph of the cross-section of one of the extrudates, and Figure 11B shows the corresponding coke band intensity measured in the indicated area. Clearly this sample shows no large-scale variations in coke content, but some small-scale deviations can be observed. These coincide with the presence of dark spots or inclusions mentioned before. Once we had established that these small-scale deviations hardly play a role in the global coke distribution of the pellets, coke profiles were measured at lower spatial resolution (50 µm) and by performing line scans in stead of area scans, to speed up the analysis procedure.
Using the line scan procedure, several cross-sections of each sample were analyzed for coke content. Figures 12 through 14 show the extremes of the observed coke profiles for each spent catalyst in series B through D. The line scans of catalysts A1 and A2 (not shown) from the pilot tests confirmed that the coke distribution of these samples is homogeneous within limits of
A B C
Figure 10: Burning of coke by laser irradiation: (A) Microscope image before mapping, (B) Raman map of the same area (dark color indicates more coke) and (C) magnified image of the large dark spot in figure A after mapping.
Catalysts B1 and B2, which were retrieved from a commercial reactor after premature shutdown, both show essentially homogeneous coke profiles (Figure 12). Also the total amount of coke appears to be quite low, looking at the total Raman intensity, in spite of the fact that the reactor of sample B2 was shut down because of fire in the unit.
As can be seen in Figure 13, the coke distribution of catalysts C1 and C2 is also essentially homogeneous, but the Raman signals for coke are significantly higher. These samples were collected at the end of run in a commercial hydrotreater processing a heavy feed, which can explain the apparent higher coke content.
A
B
Figure 11: Microscope photograph of spent catalyst A1 (A). The rectangle indicates the mapped area for coke analysis shown in (B). Dark color indicates more coke.
edge center edge
Coke signal (a.u.)
edge center edge
Coke signal (a.u.)
B1 B2
Catalysts D1 and D2, which were applied as a guard bed in an industrial reactor, clearly show a different picture (Figure 14): the observed coke profiles are highly nonhomogeneous, and more coke is observed near the center of the extrudate. In some particles typical “M-shaped” profiles can be seen. The wide-pore catalyst (B2) that was unloaded from the same reactor exhibits similar profiles as the standard catalyst (B1), albeit less pronounced.
Discussion
SEM/EDX
The SEM micrographs show that the morphology of the catalyst extrudates studied in this investigation is essentially homogeneous apart from several small “inclusions” in the alumina matrix. Other researchers observed similar inclusions in different catalyst pellets [25]; apparently this is a general feature of extruded catalysts. From the visual appearance of the inclusions it seem
edge center edge
C o ke s ignal (a .u .)
edge center edge
Coke signal (a.u.)
C1 C2
Figure 13: Coke profiles of spent catalysts C1 (end of run) and C2 (same, but wide-pore).
edge center edge
Coke
signal (a.u.)
edge center edge
Coke
signal (a.u.)
D1 D2
cracks on the surface, and in the case of sample A2, three-dimensional protrusions on the surface of the cross-section were observed with SEM. These extraordinary artifacts are most probably caused by heavy oil fractions adsorbed inside the pellet, which were forced to the clean surface of the cross-section by the vacuum applied during the gold sputtering process. Also the high fluorescence in the preliminary Raman spectra indicates that these catalyst pellets, which appeared to be dry, still contained some oil residue.
EDX showed that the active metals were homogeneously distributed throughout the catalyst pellet. Because the catalysts of all sample series were prepared using a similar procedure, we can assume that in every sample the active metals are evenly distributed. Hence, we conclude that the observed coke deposition profiles were not caused by a maldistribution of the active metals.
Raman
An elaborate Raman study on the sulfiding of molybdenum-based catalysts was done by Schrader and Cheng [26, 27]. For the unsulfided (oxidic) system they observed Raman bands around 220, 350 and 580 cm-1, and a very strong band at 945 cm-1 with a shoulder at 918 cm-1.
These wavenumbers correspond very well with the Raman spectrum of the oxidic catalyst precursor shown in Figure 7, and are attributed to polymolybdate species. After sulfiding, Schrader and Cheng [26] observed two very sharp Raman bands at 380 and 405 cm-1. These bands correspond to the vibration of MoS2 layers, and can clearly be identified in the Raman spectra of the sulfided and
spent catalyst shown in Figure 9. These results suggest that the MoS2 surface structure is stable,
even after exposure to air.
In the Raman spectra of the spent catalyst we can clearly see the D and G band of the deposited coke at 1340 cm-1 and 1604 cm-1 respectively. Both bands are very broad, which indicates that the coke is highly inhomogeneous in nature, which makes it very difficult to characterize using Raman. For reasons already mentioned in the introduction, the Raman spectra were not deconvoluted; only the total integrated peak area was used for further evaluation.
Coke profiles
By comparing the coke deposition profiles of the different catalyst samples used in this study, we can conclude that significant changes in coke distribution in the catalyst pellet only appear after a long time on stream. Both the catalysts from the laboratory tests (series A) and those from a short run in a refinery hydroprocessing unit (series B) show a low amount of coke with a homogeneous distribution. Of course, comparison of the total amount of coke based on the Raman intensities only gives a rough estimation, for reasons explained in the introduction. This estimation is justified by the fact that spent catalysts B1 and B2 have similar Raman intensities in the coke region, which corresponds to their almost identical coke content (about 15 wt% C), as reported in a previous study by Eijsbouts and Van Leerdam [28]. They found that the main difference between these samples was the structure of the active phase; the catalyst from the emergency shutdown had suffered from extreme sintering of the MoS2 crystallites. Remarkably, this had no measurable effect on the coke
content. At the end of run in an industrial hydrotreater (for catalyst series C after one year on stream) the amount of coke appears to be significantly higher, which is of course common for these processes [29]. However, no distinct profiles can be observed in the coke distribution.
In contrast, the catalysts in series D have developed very distinct coke profiles. In the wide-pore catalyst these are less pronounced than in the standard catalyst, which indicates that the development of the profiles is related to diffusion limitations inside the pores of the catalyst. It was already confirmed by other researchers that intra-particle limitations can occur in hydroprocessing, especially for (partially) deactivated catalysts [3, 5, 6, 30]. The observed M-shape (Figure 14) is typical for a sequential deactivation mechanism, where a reaction intermediate or product is a coke-precursor. Furthermore, this type of coke profile only develops in the case of strong diffusion limitations [31, 32]. The reason for this behavior in this particular case (series D) may be the application as a guard bed, the properties of the feedstock, of a combination of the two. Guard beds are usually applied to capture metal-containing compounds and other refractory components that may deactivate the catalyst downstream. The guard bed is positioned at the inlet of the reactor and will therefore be exposed to the highest levels of these deactivating compounds. In our case the feed was CGO, which is not a very heavy product, but generally contains a relatively high amount of (poly-) aromatic compounds [33]. These compounds may have difficulties diffusing through the pores of the catalyst and exhibit a strong tendency to form coke.
Analogously, Fozard et al. observed very similar coke profiles in commercial NiMo catalyst pellets after 2500 h hydroprocessing of atmospheric residue in a bench scale reactor [9]. Although atmospheric residue generally contains less aromatics, it is heavier than CGO. We can conclude that under certain conditions near the end of run in a commercial hydrotreater the reaction may be seriously hindered by intra-particle diffusion limitations, and the catalyst effectiveness factor may be reduced significantly.
Conclusions
Raman spectroscopy proved to be a very effective tool to measure coke profiles in spent hydroprocessing catalysts. It was however necessary to remove oil residue by extraction with toluene before analysis, and to protect the sample in an inert atmosphere during laser irradiation to prevent degradation of the coke deposits. The bulk structure of the MoS2 phase was found to be
stable, even after exposure to air. Qualitatively, the level of coke accumulated on the catalysts increased with time on stream in the process.
Two catalyst samples, taken from a guard bed, exhibited distinct M-shaped profiles, suggesting that coke was deposited by a sequential deactivation mechanism under a strong diffusion resistance. This means that, when the catalyst is exposed to heavy molecules for a long period of time, coke deposition can significantly reduce its effectiveness factor.
Acknowledgement
Elke Fakkeldij and Edwin Ariëns are gratefully acknowledged for performing the EDX analyses.
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Simulation of coke and metal deposition in catalyst pellets
using a non-steady state fixed bed reactor model
A mathematical model is presented, to describe the formation of coke and metal deposition profiles in catalyst pellets as a function of position in the reactor, and to predict the catalyst deactivation behavior due to pore blocking. The blocking of the catalyst pore structure was modeled using a tesselation process, described by percolation theory. The model can qualitatively describe the coke and metal deposition profiles found in hydroprocessing catalyst pellets, and the accelerated deactivation observed near the end of the catalyst life. A lab scale HDM experiment was simulated as a case study: The agreement between the modeled and experimental deposition profiles is very good for the top section of the catalyst bed, but less good for the bottom section. For a good quantitative description, detailed information on the deactivation mechanism and kinetics are required.
Introduction
Catalyst deactivation is a crucial factor in the design and operation of catalytic processes. In hydroprocessing, the three main causes for deactivation are (1) the poisoning of the catalyst’s active sites by strongly adsorbing molecules like nitrogen compounds or coke, (2) the blocking of the pore structure by the deposition of coke or metal sulfides (also called fouling), and (3) sintering of the active phase. The rate of deactivation depends on the process conditions, the feed, the catalyst itself, and the deactivation mechanism. For instance, increasing the hydrogen pressure will suppress coke deposition but accelerate the deposition of metals. Heavy feeds will in general cause more coking, and contain more metals [1].
In hydroprocessing, catalysts show a fast deactivation during the first few days on stream, after which a long period of almost stable activity is attained (see Figure 1). At the end of the run, which can be up to two years of operation, the activity drops again fast and the catalyst needs to be regenerated [2]. The so-called start of run deactivation is due to a rapid initial coke deposition process and the adsorption of poisons from the feed. This process is mainly controlled by the type of the catalyst, e.g. the stability of the active sites or the acidity of the support, and by the feed composition. The second deactivation process is mainly due to the blocking of the pore structure by deposited metals and coke, and partly to a loss of dispersion [3]. In this case the pore structure of the catalyst plays a major role, especially near the end of the catalyst life.
The aim of the present work is to derive a model that can describe the deposition of metals or coke and resulting blockage of the catalyst pore structure on the reactor level, using percolation theory. In particular, we will focus on the development of deposition profiles in the catalyst pellets. Depending on the diffusion resistance and deposition kinetics, different profiles will be found, as shown in Figure 2. Modeling of this process will provide a better insight in the deactivation mechanism of hydroprocessing catalysts, and can be used to predict their deactivation behavior in an industrial reactor. Case studies will be presented to compare the modeling results to experimental
activity metals coke 0 25 20 15 10 5 D ep os its (w t% ) Relat iv e ac tiv ity 0 200 300
Time on stream (days) 100
0 1
