Fluid catalytic cracking: Feedstocks and reaction mechanism

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(1)Fluid Catalytic Cracking Feedstocks and Reaction Mechanism.

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(3) Fluid Catalytic Cracking Feedstocks and Reaction Mechanism. 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 vrijdag 3 maart 2006 om 10:30 uur door Xander DUPAIN scheikundig ingenieur geboren te Ridderkerk.

(4) Dit proefschrift is goedgekeurd door de promotor: Prof. dr. J.A. Moulijn Toegevoegd promotor: Dr. ir. M. Makkee Samenstelling promotiecommissie: Rector Magnificus, Prof. dr. J.A. Moulijn, Dr. ir. M. Makkee, Prof. dr. W. Buijs, Prof. dr. ir. G.B. Marin, Prof. dr. R.A. van Santen, Dr. C.J. Schaverien, Dr. P. O’Connor, Prof. dr. ir. M.-O. Coppens. voorzitter Technische Universiteit Delft, promotor Technische Universiteit Delft, toegevoegd promotor Technische Universiteit Delft Universiteit Gent Technische Universiteit Eindhoven Shell Global Solutions International B.V. Albemarle Catalysts Company B.V. Technische Universiteit Delft (reservelid). This research was financially supported by: Chapter 5-10: Chapter 3:. Chapter 4:. Shell Global Solutions International B.V. Shell Global Solutions International B.V. Albemarle Catalysts Company B.V (formerly AKZO Nobel Catalysts B.V.) Netherlands Organisation for Scientific Research (NWO) Engelhard Corporation Netherlands Organisation for Scientific Research (NWO). ISBN-10: 90-9020417-2 ISBN-13: 978-90-9020417-8 Printed by Ponsen & Looijen, Wageningen Cover design: Hanneke Schreurs Copyright © 2006, X. Dupain.

(5) Table of contents Chapter 1:. Introduction. 1. Chapter 2:. Experimental set-up. 21. Chapter 3:. Cracking behaviour of organic sulfur compounds under realistic FCC conditions. 31. Chapter 4:. Aromatic gas oil cracking under realistic FCC conditions. 51. Chapter 5:. Production of clean transportation fuels and lower olefins from Fischer-Tropsch Synthesis waxes under FCC conditions. 71. Chapter 6:. Cracking of a rapeseed vegetable oil under realistic FCC conditions. 103. Chapter 7:. Optimal conditions in fluid catalytic cracking; a mechanistic approach. 135. Chapter 8:. Reaction mechanism for the catalytic cracking of paraffins. 175. Chapter 9:. Thermodynamics and surface alkoxide complexes in fluid catalytic cracking. 215. Chapter 10: Application of comprehensive two-dimensional gas chromatography in a mechanistic study on fluid catalytic cracking. 257. Chapter 11: Evaluation. 277. Appendix:. 295. GCxGC chromatograms. Samenvatting. 301. Publications & presentations. 305. Dankwoord. 309. Curriculum Vitae. 311.

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(7) 1 Introduction. Abstract An introduction to refining and Fluid Catalytic Cracking (FCC) is presented in this chapter. The major refinery operations and quality of gasoline and diesel fuels are discussed. An overview of a modern FCC unit and FCC catalysts is given; Process operation and the composition of cracking catalyst are discussed. The general trends and developments in refining and FCC are briefly highlighted. Finally, the outline of this thesis is presented.. 1.

(8) 1.. Oil refining. 1.1. Introduction. The global demand for petroleum products will grow for many years to come and, as a consequence, the capacities of the refining and petrochemical industries will increase. The purpose of oil refining is the production of fuels for transportation, power generation, heating purposes, and the production of base chemicals [1]. Crude oil consists of a wide range of hydrocarbons. These hydrocarbons can roughly be divided into different groups according to their molar mass and structure [2,3]. The most prominent are the groups of the paraffins (linear and branched), the naphthenes (mainly unsaturated five and six-membered carbon rings) and aromatics (benzene derivatives and poly-aromatics). Furthermore, a large amount of molecules are present that contain metals and one or more heteroatoms, e.g. sulfur, nitrogen, oxygen, or metals. In figure 1 a simplified overview of a modern refinery is given. The incoming crude oil is fractionated into different sub-streams by atmospheric- and vacuum distillation. In order to meet the market demand and to use oil as transportation fuel the size of the molecules needs to be reduced and the hydrogen/carbon (H/C ratio) to be increased. Fuel Gas LPG. LPG. Alkylation. Alkylate LPG. Straight-run Gasoline Naphtha. Catalytic Reforming. Hydrotreating. Middle Distillates Atmospheric Gas Oil. Hydrotreating. FCC. Reformate. Gasoline. Hydrotreating. Solvents. Gasoline Kerosene Slurry Oil. Cycle Oil. Vacuum Distillation. Vacuum Gas Oil. Hydrocracking Lube Basestocks. LPG & Gas Gasoline, Naphtha, Middle Distillates Hydrowax. Solvent Extraction. Lube Oils Solvent Dewaxing. Treating & Blending. Crude Oil. Atmospheric Distillation. LPG & Gas. Diesel. Heating Oil. Lube Oil. Waxes Vacuum Residue. Propane Deasphalting. Greases LPG & Gas Asphalt. Vacuum Residue. Visbreaking & Coking. Gasoline, Naphtha, Middle Distillates. Industrial Fuels. Fuel Oil, Asphalt Coke. Coke. Figure 1. Schematic overview of a modern-day refinery [1]. Grey blocks indicate catalytic processes.. 2.

(9) Table 1. Octane numbers of different refinery gasoline streams. Gasoline stream Light Straight-run Naphtha FCC Light Gasoline FCC Heavy Gasoline Reformate Alkylate. RON. MON. 68 93 95 99 95. 67 82 85 88 92. RON= research octane number; MON= motor octane number. Therefore, the streams from the atmospheric- and vacuum fractionation are upgraded by catalytic and thermal processes. In practice the conversion processes in the refinery operate through the “hydrogen in” or “carbon out” principle. In the first case, hydrogen atoms are added to the molecule. A typical example is Hydrocracking. In the second case, carbon atoms are removed from the molecule. Generally, thermal processes operate through the “carbon out” principle, but also Fluid Catalytic cracking is a clear example. 1.2. Quality of gasoline and diesel. The quality of the produced gasoline is defined by the octane number and is an indication for the resistance of a motor fuel to knock [4]. By definition, the octane numbers are based on a scale on which 2,2,4-trimethyl-pentane (iso-octane) is 100 and n-heptane is 0. A gasoline with an octane number of 90 has the same knock characteristics as a mixture of 90% iso-octane and 10% n-heptane. Modern gasoline-powered vehicles require a minimum research octane number (RON) of 92, 95, or 98 [1]. Branched hydrocarbons are excellent octane enhancers, and also aromatics and naphthenic compounds are beneficial. Linear paraffins possess a low octane number and are, hence, not desirable in gasoline [5]. In table 1 an overview has been given on the octane numbers of different gasoline products. Often methyl-tert-butyl ether (MTBE; 118 RON), ethyl-tert-butyl ether (ETBE; 118 RON), and (bio)ethanol (111 RON) are added to enhance the octane number even further [1,6]. Moreover, MTBE helps to reduce the carbon monoxide emissions [7]. Opposed to gasoline, linear paraffins are highly desirable for diesel fuels, whereas aromatics have an adverse effect on the quality. The main quality index of diesel is defined by the cetane number. By definition, n-hexedecane (cetane) has a value of 100 and α-methyl-naphthalene has a value 0 [1]. In table 2 an overview of the cetane numbers of different refinery gas oil stream is presented.. 3.

(10) Table 2. Cetane numbers of different refinery gas oil streams. Gas oil stream Straight-run gas oil FCC cycle oil Hydrocracker gas oil Thermal gas oil. Cetane number 40-50 0-25 55-60 30-50. The gasoline octane number and diesel cetane number are the main quality indices, but certainly not the only measures. For instance, the transportation fuels have to apply in different geographical areas under different conditions and, therefore, their composition will be variable. All the different products from the refinery have to meet strict quality specifications imposed by legalisation [8-11]. Catalysts used in automotive after-treatment devices are particularly sensitive for the poisoning by the sulfur in the transportation fuel. Moreover, sulfur and nitrogen form oxides upon combustion that are harmful to the environment. The same applies for aromatics and olefins that can form harmful species. Recently it has become clear that sulfur in diesel directly contributes to particle emission [8,11]. 1.3. Main refinery operations. Straight-run naphtha from the distillation towers has typically a low octane number, which makes it less attractive as a fuel. In order to increase the octane number the naphtha is subjected to catalytic reforming and blended with gasoline from the FCC unit [7,12]. The aromatic- and olefin content are enhanced over a platinum-chlorine catalyst. Moreover, some skeletal rearrangements take place. The reforming process has as an additional advantage that it produces valuable hydrogen as a by-product, which can be used elsewhere in the refinery, for instance in hydrotreating and hydrocracking units. The propene and butenes from a part of the LPG (C3-C4) in the refinery are alkylated with ibutane to alkylate gasoline with an octane number of more than 95. Although the produced gasoline components as such are desirable from an environmental point of view the process is environmentally unfriendly as it commonly takes place in the presence of a liquid acid catalyst, typically H2SO4 or HF. A recent trend in alkylation is the development of solid acidic catalysis, eliminating the main disadvantages of the current processes [13-15]. The down-stream catalysts used in the refinery processes and the catalysts of automotive after-treatment devices are particularly sensitive for poisoning by the hetero-atoms in. 4.

(11) hydrocarbon molecules, e.g. sulfur, nitrogen, oxygen, and metals. These hetero-atoms can be catalytically removed by catalytic hydrotreating; in this process transformation take place into gaseous molecules (H2S, NH3, H2O) that are easily separated from the liquid product stream. Typically trickle-beds of CoS/MoS or NiS/MoS catalysts are used in which the liquid oil is contacted at high pressure with hydrogen. The heart of the refinery consists of a Catalytic Cracking unit and a Hydrocracker unit. In both units the size of the molecules is reduced. Fluid Catalytic Cracking is a ‘carbon out’ process. Its purpose is to upgrade heavy oil fractions through catalytic cracking into valuable products such as gasoline, Light Cycle Oil (LCO, a diesel-blending component), and lower olefins (C3= and C4=). This so-called ‘whitening the barrel’ is depicted in figure 2 for the approximate product ranges. During the cracking reactions over an acidic zeolite Y coke is deposited on the catalysts. The main products from FCC is gasoline that posses a relatively high-octane number (> 90). Moreover, it produces large amounts of propene, which is a base chemical, and butenes that can be used in the alkylation process. Isobutene can be reacted with methanol or ethanol to produce high-octane number methyl-tert-butyl ether (MTBE) and ethyl-tert-butyl ether (ETBE), respectively. The bottom product form the FCC unit, heavy cycle oil, is usually recycled or upgraded in a Hydrocracker unit. Hydrocracking is a ‘hydrogen in’ process, and practically operates similar to the hydrotreating process with the difference that an acidic function has been added to the catalyst and the conditions are more severe. Due to the hydrogenation the products, mainly paraffins that lie in the diesel-range, are saturated. The heavy bottoms (Hydrowax) can be either used as waxes or upgraded in the FCC unit to produce high-octane gasoline.. LCO (215-325 oC). Dry Gas (H2,C1-C2) LPG (C3-C4). Gasoline (C5-215 oC). LCO (215-325 oC) HCO (325+ oC) FCC Feed. HCO (325+ oC) FCC Product. Figure 2. Whitening of the barrel by FCC. The data are typical for the FCC process.. 5.

(12) Due to the presence of hydrogen coking is a less problem than that in FCC and the process operates generally in trickle-bed reactors. The residue from the vacuum distillation unit consists of high-boiling hydrocarbons. It is generally highly-aromatic and, hence, easily forms coke in both catalytic and thermal upgrading processes. In propane deasphalting the feed is brought into contact with propane in which the paraffinic fraction dissolves, but the heavy aromatic don’t. The propane is removed by evaporation and recycled. In solvent dewaxing the linear paraffins are separated from the branched paraffins in the reaction mixture. Branched paraffins have a lower crystallisation temperature and are, hence, more suitable as lubricants than the linear paraffins. Visbreaking and coking are processes that convert the heaviest fraction from the distillation. The ExxonMobil Flexicoking (carbon out) process and Shell HYCON (hydrogen in) are sophisticated processes to obtain as much useable products as possible from this fraction. The heaviest products are bitumen and solid coke. The former is used as asphalt, whereas the latter can be used as product for semi-conductor industries, depending on its quality.. 2.. Fluid Catalytic Cracking. Fluid Catalytic Cracking (FCC) is one of the key processes in the modern refinery and one of the largest applications of catalysis with a worldwide processing capacity that exceeds 500 million tons of oil per year [1]. The conversion of the oil feed takes place in an entrained-flow reactor (riser reactor) in the presence of a solid acidic catalyst, as shown in figure 3 [1,16]. The endothermic cracking reactions proceed over a catalyst that commonly consists of highly-acidic zeolite crystals (210 µm) dispersed in a moderately-acidic silica-alumina matrix (50-70 µm) [1]. The catalyst rapidly deactivates due to the deposition of coke. In the top of the riser reactor the catalyst is separated from the gaseous products by cyclones and stripped from adsorbed hydrocarbons with steam. Subsequently, the coked catalyst is led to the fluidised-bed regenerator reactor where the deposited coke is combusted and it regains activity. The system operates in a heatbalance mode [1,17-21]. The feed is generally mixed with steam to facilitate the atomisation of the oil and smooth catalyst flow [22]. In figure 4 a flow scheme of a modern FCC unit with product separation section is shown.. 6.

(13) Reactor. Regenerator 2-Stage Cyclones. Flue Gas. 2-Stage Cyclones. Coked Catalyst. Stripping Steam. 700 oC 2 bar. Riser Air. 525 oC 1 bar. Regenerated Catalyst. Feed. Figure 3. Schematic overview of a commercial Fluid Catalytic Cracking (FCC) unit. Flue Gas Waste Heat Boiler. Dry Gas. G/L Sep.. Propane/Propene Cyclones L/L Sep.. Water. Catalyst Fines. Coked Catalyst. Steam. Reg. Catalyst. Riser Butanes/Butenes. Expansion. Air. Gasoline. Compression. Light Cycle Oil Steam. Heavy Cycle Oil. Feed. Slurry Oil Regenerator. Reactor. Fractionator. Absorber. Debutaniser. Depropaniser. Figure 4. An FCC unit with separation train.. 7.

(14) 3.. Cracking catalysts. Modern cracking catalysts are generally composed of faujasite (zeolite Y) crystals held together by a silica-alumina matrix. These catalysts can be produced by addition of an acidic clay (usually kaolin platelets) and acidic faujasite zeolite crystals to and silica-alumina slurry. The resulting mixture is spray-dried to form microspheres of 50-70 µm in diameter that after washing, drying, and calcination yield the desired catalyst [1]. The incorporated zeolite crystals have a size of 2-10 µm and are highly-active as a result of their high acidity. Moreover, they posses a high selectivity to the desired products, high thermal stability, and low tendency to produce coke. 3.1. Zeolites. Zeolites are crystalline alumina-silicates, which are structurally unique in having regular cavities and pores with molecular dimensions. As a result of the small pore sizes they possess a high molecular shape selectivity and a high surface area [23]. The regular pattern enables the zeolite to act as a selective sieve for a number of compounds. This shape selectivity makes it possible that only certain molecules can reach the catalytic sites (reactant selectivity), certain products can leave the catalytic sites (product selectivity) and/or certain intermediates can be formed (transition state selectivity). Different types of zeolites are used for a wide variety of chemical processes [24]. In figure 5 the most common zeolite faujasite Y used in Fluid Catalytic Cracking is depicted. The primary building blocks for faujasite Y consists of 24 SiO4 and AlO4 tetrahedra, forming a truncated octahedron or sodalite cage. The sodalite cages are arranged in such an array that it is connected to four other sodalite cages, linked by six bridging oxygen ions.. (a) Sodalite cage. (b) Faujasite structure. Figure 5. Structure of a faujasite zeolite [21].. 8.

(15) 2. Si. O. Al. H O. Si. Brønsted acid site. Si. O. Al. +. Si. +. Si. O. Al. O. Si. + H2O. Lewis acid site Figure 6. Brønsted and Lewis acidity in zeolites [21].. The apertures are 0.74 nm and are surrounded by a twelve-membered oxygen ring. The supercage in the centre of this structure is surrounded by ten sodalite units and is large enough to contain a 1.2 nm sphere. The sodalite cages themselves have six-membered ring entrances with a diameter of 0.24 nm [21,25]. The acidity of zeolites is based that silica from the lattice is replaced with alumina, as represented in figure 6. The lower charge of the alumina atom (+3) makes that an additional cation is required to maintain electronic neutrality. The use of a proton as additional cation gives the aluminosilicate its acidic character and represent its Brønsted acidity. The location of protons is, therefore, very important to describe the activity. When the density of proton donor groups in zeolites is low than the proton donor strength is high. Evidently the proton donor strength is greatest of groups that are associated with AlO4- tetrahedra having the smallest number of Al neighbours. When the hydrogen-form zeolite is heated to a high temperature, a water molecule is released leaving an unsaturated Al3+ ion (Lewis acid site) [21,22,25]. The zeolite pores have a solvent-like character and enables high hydrogen-ion transfers (hydrides and protons) [21]. For this reason zeolites are more active than the amorphous alumino-silicates, which lack the small pores. 3.2. Matrix. The FCC zeolite possesses an extremely high acidity. It cannot be used as cracking catalyst as such, not in the least due that it would be subject to severe mass-transfer limitations imposed through its microporous structure. The large feed molecules would be unable to diffuse rapidly into the zeolite channels. Moreover, the catalyst particle requires a defined size to enable satisfactorily fluidisation characteristics in the riser and regenerator. Therefore, the zeolite crystals are dispersed in a matrix that posses a mesoporous structure. Usually silicaalumina gels are used that posses a medium acidity and pre-cracks the largest molecules into smaller hydrocarbons that can enter the zeolite. Fillers, commonly kaolin clays, are added to dilute the activity of the catalyst. Binders (kaolin, bentonite) serve as a glue to hold the ingredients together [1,18,19,26,27].. 9.

(16) Matrix. Zeolite (2-10 µm). 50-70 µm Figure 7. Composition of an FCC catalyst, after ref. [1].. In figure 7 a schematic representation of a catalyst particle is given. These are not the only purposes of the matrix. The matrix protects the active zeolite content for the attrition that the particles experience with each other and the equipment hardware. In fact, about 5 % of fresh catalyst is added every day to compensate for the loss that is caused by the attrition that results in fines and deactivation [1]. Since the catalyst particles act as heat carrier from the regenerator to the riser the matrix should possess sufficient heat capacity properties. Additionally it acts as a sink for sodium ions and other contaminants. 3.3. Diffusion in FCC catalysts. In most cases the diffusion of reactants in the matrix lies in the so-called Knudsen regime. The pores (mesopores) of the matrix significantly influence the movement of the diffusing molecules. The pores of the zeolites are generally an order of magnitude smaller, i.e. micropores) than the pores of the matrix and the diffusion belongs to the ‘configurational regime’. The size of the molecules is in the order of the channel dimensions and molecules may be unable to pass each other [23,28]. 3.4. Irreversible deactivation. Catalyst deactivation is a common problem in FCC. In the riser the catalyst is deactivated through the deposition of coke. This reversible type of deactivation is eliminated through the combustion of these carbonaceous species in the regenerator. More of a problem is the variety of physical and chemical changes that the catalyst undergoes, denoted as irreversible deactivation. Under the severe conditions in the regenerator, the zeolite component of the catalyst is heavily dealuminated and decomposed and shows a strongly reduced in activity [25]. Moreover, a collapse of the matrix structure can alter the cracking activity and accessibility of. 10.

(17) the catalyst [29]. Another problem is the deposition of contaminant metals (vanadium, nickel, iron, etc.) from the feed that alter the catalyst properties. Nickel deposits throughout the catalyst particle and enhances (de)hydrogenation reactions, leading amongst other to increased coke formation. Vanadium and iron are less mobile than nickel and deposits more on the outer shell of the particle. It mainly harms the zeolite framework, thereby, causing loss in activity [7,30]. Due to the fact that every day a fraction of the catalyst inventory of an FCC unit is replaced an age distribution is created. The added fresh high-active catalyst undergoes a decay process as a result of the ageing, (hydro)thermal in an FCC unit, and deposition of contaminant metals from the feed until it reaches an apparent steady-state activity. The FCC catalyst inventory is commonly denoted as equilibrium catalyst or shortly as E-cat. 3.5. Modification of zeolite Y. Increased conversion rate is generally acquired through an increase in the number, strength, or stability of the acid sites in the zeolite and/or matrix. The surface area, level of ultrastabilisation, and cation exchange are used to modify the zeolite characteristics. The ‘standard’ acidic zeolite HY is often modified to enhance certain properties. Dealumination of zeolites is often applied to increase the SiO2/Al2O3 ratio that adds thermal and hydrothermal stability to the zeolite [24]. These catalysts contain lower framework alumina contents and are denoted as ultra-stable Y (USY). Due to the dealumination the concentration of acid sites obviously decreases and a lower intrinsic activity is the result [27]. On the other hand, the decrease of acid site density will lead to an increased acidity and are suitable to crack refractory feeds, also because during the ultra-stabilisation procedure mesopores are formed that are beneficial for the diffusion of larger molecules [27]. Moreover, these zeolites tend to produce higher amounts lower olefins at the expense of gasoline [31]. A typical SiO2/Al2O3 ratio lies around 50. The cation exchange of zeolite Y with rare earth (RE) metals (lanthanum, cerium, praseodymium, neodymium, etc.) leads to a stronger acid site strength, but will also affect other reactions such as hydrogen transfer activity [24,27]. Coke deposition increases with the RE content [27]. Rare earth ion-exchanged faujasites (RE-HY) find considerable technical importance due to their superior catalytic properties. Rare earth elements can prevent aluminium losses form the Y-zeolite structure and enhance the structural resistance to severe hydrothermal conditions in the FCC process [27,32,33]. As a consequence, these catalysts are only slightly dealuminated and typically have a SiO2/Al2O3 ratio of about 15.. 11.

(18) The ultra-stabilisation of the zeolite (RE-HY) can also lead to an even lower hydrothermal deactivation rate. Such RE-USY catalysts have a lower number of acid sites as compared to RE-HY catalysts due to the dealumination of the zeolite framework. Overall, the choice of the zeolite type in the catalyst particle in relation with the design of the (support) matrix depends on the process conditions, feed characteristics, and desired product spectrum. The activity and selectivity of zeolites can be reasonably correlated to the unit cell size (UCS) that is a measure for the silica/alumina ratio of the zeolite [34]. Zeolites with a low UCS have a high silica/alumina ratio and, as a consequence, a reduced number of acidic sites. The content of HY or RE-HY zeolite in the catalyst matrix is typically 10-30 wt%. In order to compensate for the lower intrinsic activity of the USY or RE-USY zeolite the content of these zeolites in the catalyst particle is usually higher, viz. 30-40 wt% [35]. The acid site strength and density of the matrix can be modified by changing the type of constituents, e.g. alumina and silica. 3.6. Catalyst additives. Additives are often added to FCC catalysts in order to manipulate certain features. For instance, ZSM-5 additives that are composed of acidic ZSM-5 zeolite crystals in a silicaalumina matrix are added to enhance the cracking of linear and mono-branched paraffins and olefins from the gasoline fraction to valuable lower olefins (C3 and C4). The cracking of the linear hydrocarbons from the gasoline fraction obviously leads to a lower yield. The selective removal of linear hydrocarbons results in a higher aromatic content in the gasoline and, hence, an increased octane number [27,36]. The sulfur of the feed that ends up in the coke fraction on the catalyst can be removed through cerium and magnesium additives that promote oxidation of SO2 to SO3 in the regenerator and forms sulphate complexes with the metal [1,37,38]. In the reducing atmosphere of the riser reactor and stripper SOx is converted to H2S that leaves the system with the product and can be recovered as elemental sulfur by Claus plants. Another example of additives are combustion promoters (e.g. platinum, palladium) that catalyse the full oxidation of coke to CO2 and limit the formation of CO. Also additives such as vanadium traps (e.g. magnesium oxide and titanates) and nickel passivators (e.g. antimony) will reduce the effects that contaminant metals (e.g. vanadium and nickel) have on the active zeolite component of the catalyst and on its cracking selectivities.. 12.

(19) 4. Trends and developments in refining and Fluid Catalytic Cracking. 4.1. Refining. The world production of liquid hydrocarbons is expected to reach a maximum in about 20152035, and depending on the growth scenarios the ultimate reserves should cover only 60 to 100 years [39]. Although the present petroleum sources are sufficient to accommodate the demand for the first half of the 21st century, the wide distribution of locations poses a problem as well. The major part of the proven petroleum reserves is located in geographical zone where the political climate is unsettled [39]. Additionally, most consumer zones are located far from the production zone, which calls for intensive transportation activity. Worldwide, oil companies are extensively investigating the opportunities of other sources of energy. Solar power and wind power are clean alternative sources, but will not easily fulfil the demand of energy for an ever-growing world. A key issue is to promote the growth of development of third world countries while controlling pollution and consumptions [39]. Coal, natural gas, and shale oil are potential sources for the production of liquid hydrocarbons [13]. The gasification of these resources to synthesis gas, followed by Fischer-Tropsch Synthesis has already proven to be a promising application for the production of mainly diesel fuels [7,40,41]. Tar sands are another option. These resources contain relatively heavy hydrocarbons, which for upgrading need to be extracted and adapted for refining. Although these three examples are good alternatives for oil they also belong to the group of limited resources and, in that sense, are not sustainable. A more urgent problem is that their use disturbs the environment. Bio-oil derived from biomass can be better alternatives for the long term. Bio-ethanol and vegetable oils can be derived from crops and can be upgraded to fuel applications. Another option is the pyrolysis of biomass to so-called pyrolysis oil that potentially can be converted in useful products. In line with coal and natural gas gasification of biomass to synthesis gas and subsequently to transportation fuels seems a logical manufacturing route [13]. One of the prospective energy carriers is hydrogen that is destined to play an important role as energy carrier in the future [13]. An example related to this source is the fuel cell technology. With this technology hydrogen can be more efficiently used as energy provider than conventional hydrocarbons and emits as a combustion product only water. Though, it should be recognised that for the (cost) efficient production of hydrogen the major source are hydrocarbons.. 13.

(20) Irrespective of the above-discussed alternatives it is evident that refining will continue to play an important role in society, but will be subject to dramatic changes in the coming decades. Refining has always been a dynamic business that continuously changes to meet society’s demands. For example, in the past thirty years refining has become more complex in order to fulfil the demands in quantity and quality of products. These included increased conversion capacity from heavier oil fractions (mainly FCC and Hydrocracking), reduction in sulfur (Hydrotreating), enhanced gasoline production (Catalytic Reforming, Alkylation), and reduction of pollutants (NOx, SOx, VOCs, greenhouse gasses, and contaminated liquids and solids) [7,37]. Biorefining is a more sustainable option. 4.2. Fluid Catalytic Cracking. FCC operation is subject to constant modifications because of product quality requirements (particularly in automotive fuels) and differences in feed composition [42]. FCC gasoline has a high content of aromatics, olefins, and sulfur. Within the next ten years both gasoline and diesel fuel will be restricted to ultra low sulfur concentrations as a result of the environmental regulations [37,43,44]. But also the hydrocarbon spectrum will shift from aromatics to naphthenes and paraffins that burn cleaner [7]. The produced FCC gasoline needs to undergo hydrodesulfurisation with minimum loss of motor octane number. In Europe the market demand is shifting more towards diesel whose application as transportation fuels is more efficient from energetical point of view. Light Cycle Oil (LCO) from catalytic cracking units is an important diesel-blending component, but contains relatively high amounts of aromatics and sulfur. It possesses a relatively low cetane number and, although hydrotreating and blending can improve this, this is quite difficult (only 6-15 points increment) [39]. Diesel is more efficiently produced by Hydrocracking this process will gain more and more importance in the refinery as compared to catalytic cracking regardless its high hydrogen demand. Although the gasoline will remain a significant market share the coming decennia and the FCC capacity will increase significantly to satisfy market demands [45]. Moreover, it is of utmost importance to evaluate the potential of catalytic cracking in applications other than traditional gasoline producer. 4.3. Lower olefins. Propene is an important base chemical for the production of plastics. Propene demand was approximately 39 million tons in 1995 and this is forecasted to increase to almost 90 million tones by 2015 [46]. Although the need for propene is growing faster than for ethene the coproduction of propene from steam crackers is expected to decline as plants are optimised to produce higher-value ethene [47]. Due to its mode of operation FCC is a major source of 14.

(21) propene (and butenes). Conventional FCC units produce about 4 wt% propene [7,48]. Adding ZSM-5 additives to the catalyst is a proven technology to enhance the production of lower olefins at the expense of gasoline [7,39,48-53]. Resid Fluid Catalytic Cracking (RFCC) is a technique that aims to convert heavy oil fraction under more severe conditions (higher temperature, higher catalyst-to-oil ratio) than conventional FCC to especially lower olefins [50,54,55]. Increasing the temperature leads to an increase in light olefins, but also produces undesired dry gas (H2, C1-C2) [50,51]. Additives and reactor configurations for the recracking of naphtha olefins to propene will also gain importance in the future [7,12,40]. The development of alternative state-of-the-art catalysts for improved lower olefins is another potential alternative [47]. A zeolite that is highly active and selective for the cracking of linear olefins and paraffins is desirable. Additionally, it should be poor catalysts for carrying out hydride transfer reactions, yet should still be able to reach high bottom conversion [51]. 4.4. Alternative catalysts. The feed sources are expected to evolve towards heavy resides that contain higher levels of metals, (poly)aromatic compounds, and heteroatom species [34,37]. Catalyst manufactures continuously strive to optimise their catalyst allow the processing of such heavy feeds. In order to increase the tolerance of FCC catalysts to metals advanced matrix and zeolite technologies are developed [34,37]. But also for the conversion of heavy resid molecules the matrix and zeolite must comply with specific characteristics, since mass-transfer will play a more role [56]. It is, therefore, important to construct a catalyst in such way that to ease the mass-transport to the zeolite crystals. There is evidence that the zeolite external area is responsible for a portion of the bottoms cracking activity of the catalyst [37]. Aromatics species and aromatic sulfur are refractory to crack and a catalyst that potentially could convert aromatics during FCC into naphthenes or paraffins by enhanced hydride transfer would be a breakthrough [43,44]. However, under FCC conditions the hydrogenation of aromatics to naphthenic compounds is thermodynamically limited [57]. Prevention of the formation of aromatics during FCC seems, therefore, of utmost importance and the hydride transfer activity of the catalyst will play a crucial role here. 4.5. Alternative (sustainable) feedstocks. Other potential options are to use renewable, more sustainable feedstocks. Fischer-Tropsch Synthesis products are highly paraffinic and extremely pure (no aromatics, sulfur, and nitrogen). Currently Fischer-Tropsch Synthesis waxes are upgraded through mild hydrocracking into clean diesel-range fuels [40]. Alternatively FCC could produce gasoline and diesel without the requirement of hydrogen. Moreover, lower olefins could be produced. 15.

(22) The category of well-defined feedstocks is also applicable to vegetable oils and fatty acid. The cracking of pyrolysis oil could be an application, but these bio-oils are generally high on aromatics and, hence, hydrogenation/hydrocracking seems more feasible for the production of fuels.. 5. Outline of this thesis. The chapters in this thesis have been written as separate publications and can be read independently. As a consequence, it cannot be avoided that some overlap between them occurs. In chapter 2 the standard experimental procedure is discussed. In this chapter the microriser set-up, analytical methods, and calculation procedures are explained. Specific information on the experimental procedures related to the individual chapters will be discussed there. In the first part of this thesis the catalytic cracking behaviour of different feedstocks and hydrocarbon classes is discussed. The catalytic cracking behaviour of organic sulfur compounds in FCC feedstocks have been elucidated in chapter 3. In chapter 4 the catalytic cracking of aromatics is described. Opposed to conventional FCC a Fischer-Tropsch Synthesis wax has a highly paraffinic nature and the catalytic cracking of this feed and the potential to use it as a feedstock for FCC is discussed in chapter 5. The catalytic cracking of a rapeseed vegetable oil and its product distribution is dealt with in chapter 6. In chapter 7 a comparison is made between different feedstocks and it is discussed how mass-transfer characteristics and process conditions affect the product distribution in FCC. In the second part of this thesis, chapters 8-10, the focus is on the mechanistic fundamentals of catalytic cracking. In chapter 8 the mechanisms for the main reactions in the catalytic cracking of paraffins are discussed. In chapter 9 the thermodynamics, mechanisms, and dynamics through surface alkoxide complexes are combined and describe the processes in FCC. In chapter 10 a demonstration is presented on how valuable kinetic and mechanistic information is derived from the catalytic cracking of a Fischer-Tropsch Synthesis wax by comprehensive two-dimensional gas chromatography. In conclusion, in chapter 11 an evaluation is made.. 16.

(23) 6. References. [1]. J.A. Moulijn, M. Makkee, and A. van Diepen, “Chemical Process Technology”, John Wiley & Sons Ltd., New York (2001) V. Simanzhenkov and R. Idem, “Crude Oil Chemistry”, Marcel Dekker, Inc., New York (2003) F.D. Mango, Org. Geochem. 26, 417-440 (1997) E.L. Marshall and K. Owen, “Motor Gasoline”, SCI Publications, Oxford (1995) J.-P. Wauquier, “Petroleum Refining Vol. 1”, IFP Publications, Éditions Technip, Paris (1995) A.G. Lucas, “Modern Petroleum Technology, Vol.2 Downstream”, 6th Ed., John Wiley & Sons, Ltd., Chichester (2000) T.G. Kaufmann, A. Kaldor, G.F. Stuntz, M.C. Kerby, and L.L. Ansell, Catal. Today 62, 77-90 (2000) G. Corro, React. Kinet. Catal. Lett. 75, 89-106 (2002) B.A.A.L. van Setten, M. Makkee, and J.A. Moulijn, Catal. Rev. 43, 489-564 (2001) R. Burch, Catal. Rev. 46, 271-333 (2004) M.S.P. Kahandawala, J.L. Graham, S.S. Sidhu, Energ. Fuel 18, 289-295 (2004) A. Corma, F.V. Melo, L. Sauvanaud, and F.J. Ortega, Appl. Catal. A. 265, 195-206 (2004) C. Marcilly, J. Catal. 216, 47-62 (2003) E.H. van Broekhoven, J. Sant, S. Zuijdendorp, and N. Winkler, International Patent WO 2005/075387, Albemarle Netherlands B.V. (2005) J.F. Himes, R.L. Mehlberg, F.-M. Nowak, “Advances in Hydrofluoric Acid Catalyzed Alkylation” NPRA Annual Meeting, San Antonio, USA (2003) M. Guisnet and S. Mignard, L’actualité chimique, Février 2000, 14-22 (2000) S. Raseev, “Thermal and Catalytic Processes in Petroleum Refining”, Marcel Dekker, Inc., New York (2003) R. Sadeghbeigi, “Fluid Catalytic Cracking Handbook”, Gulf Publishing Company, Houston (2000) J.W. Wilson, “Fluid Catalytic Cracking Technology and Operation”, Penwell Publishing Company, Tulsa (1997) P.B. Venuto, and E. T. Habib, Jr., “Fluid Catalytic Cracking with Zeolite Catalysts”, Marcel Dekker, Inc., New York, (1979) B.C. Gates, “Catalytic Chemistry”, John Wiley & Sons Inc., New York (1991) A. Corma, O. Marie, and F.J. Ortega, J. Catal. 222, 338-347 (2004) T.F. Degnan Jr., J. Catal. 216, 32-46 (2003) E. Roland and P. Kleinschmidt, “Zeolites”, Ullman’s Encyclopaedia of Industrial Chemistry, http://www.mrw.interscience.wiley.com/ueic (2000). [2] [3] [4] [5] [6] [7] [8] [9] [10] [11] [12] [13] [14] [15] [16] [17] [18] [19] [20] [21] [22] [23] [24]. 17.

(24) [25] [26] [27]. [28] [29] [30] [31] [32] [33] [34] [35] [36] [37] [38] [39] [40] [41] [42] [43] [44] [45] [46] [47] [48] [49] [50] [51]. 18. A. Brait, “Characterization and Evaluation for Fluid Catalytic Cracking”, Ph.D. thesis, Twente University of Technology, Enschede (1997) A.A. Avidan, Stud. Surf. Sci. Catal. 76, 1-39 (1993) S. Sivasanker, in ”Catalysis, principles and applications”, B. Viswanathan, S. Sivasanker, A.V. Ramaswany (Eds.), Narosa Publishing House, New Delhi, 362-373 (2002) G. Tonetto, J. Attias, and H. de Lasa, Appl. Catal. A. 270, 9-25 (2004) D.R. Rainer, E. Rautiainen, and P. Imhof, Appl. Catal. A. 249, 69-80 (2003) O. Bayraktar, and E.L. Kugler, J. Therm. Anal. Calorim. 71, 867-874 (2003) Y.G. Adewuyi, D.J. Klocke, and J.S. Buchanan, Appl. Catal. A. 131, 121-133 (1995) Z. Peiqing, W. Xiangsheng, G. Xinwen, G. Hongchen, Z. Leping, and H. Yongkang, Catal. Lett. 92, 63-68 (2004) G. de la Puente, E.F. Souza-Aguiar, F.M.Z. Zotin, V.L.D. Camorim, and U. Sedran, Appl. Catal. A. 197, 41-46 (2004) P. O’ Connor, J.P.J. Verlaan, and S.J. Yanik, Catal. Today 43, 305-313 (1998) J.S. Magee and G.E. Dolbear, “Petroleum catalysis in nontechnical language”, PennWell Publishing Company, Tulsa (1998) J. Biswas and I.E. Maxwell, Appl. Catal. 58, 19-27 (1990) R.H. Harding, A.W. Peters, and J.R.D. Nee, Appl. Catal. A. 221, 389-396 (2001) T. Myrstad, H. Engan, B. Seljestokken, and E. Rytter, Appl. Catal. 187, 207-212 (1999) C. Marcilly, Oil. Gas. Sci. Technol. 56, 499-514 (2001) Ph. Courty, P. Chaumette, and C. Raimbault, Oil. Gas. Sci. Technol. 54, 357-363 (1999) S.T. Sie, M.M.G. Senden, H.M.H. van Wechem, Catal. Today 8, 371-394 (1991) J.M. Arandes, I. Abajo, I. Fernández, M. J. Azakoiti, and J. Bilbao, Ind. Chem. Eng. Res. 39, 1917-1924 (2000) A.A. Lappas, J.A. Valla, I.A. Vasalos, C. Kuehler, J. Francis, P. O’ Connor, and N.J. Gudde, Appl. Catal. A. 262, 31-41 (2004) J.A. Valla, A.A. Lappas, I.A. Vasalos, C.W. Kuehler, and N.J. Gudde, Appl. Catal. A. 276, 75-87 (2004) R.P. Silvy, Appl. Catal. A. 261, 247-252 (2004) X. Zhao and T.G. Roberie, Ind. Eng. Chem. Res. 38, 3847-3853 (1999) P. O’ Connor, A. Hakuli, and P. Imhof, Stud. Surf. Sci. Catal. 149, 305-321 (2004) D. Decroocq, Rev. I. Fr. Petrol 52, 469-489 (1997) J.S. Buchanan, D.H. Olson, and S.E. Schramm, Appl. Catal. A. 220, 223-234 (2001) A. Aitani, T. Yoshikawa, and T. Ino, Catal. Today 60, 111-117 (2000) A. Corma, V. González-Alfaro, and A.V. Orchillés, Appl. Catal. A. 187, 245-254 (1999).

(25) [52] [53] [54] [55] [56] [57]. X. Zhao and R.H. Harding, Ind. Eng. Chem. Res. 38, 3854-3859 (1999) A.A. Lappas, C.S. Triantafillidis, Z.A. Tsagrasouli, V.A. Tsiatouras, I.A. Vasalos, and N.P. Evmiridis, Stud. Surf. Sci. Catal. 142, 807-814 (2002) A. Corma, O. Bermúdez, C. Martínez, and F.J. Ortega, Appl. Catal. 230, 111-125 (2002) M. A. Abul-Hamayel, Petrol. Sci. Technol. 20, 497-506 (2002) K.Y. Yung, P. O’ Connor, S.J. Yanik, and K. Bruno, Catalyst Courier 53, www.albemarle-catalysts.com, (2003) Chapter 9 of this thesis, “Thermodynamics and surface alkoxide complexes in Fluid Catalytic Cracking”. 19.

(26) 20.

(27) 2 Experimental set-up Laboratory once-through microriser reactor. Abstract The laboratory-scale once-through microriser reactor, used for the catalytic cracking experiments, has been described and discussed. For comparison a brief introduction has been given on the different types of laboratory-scale testing reactors for FCC. The operation procedure and analytic methods that apply for the microriser reactor have been elaborated.. 21.

(28) 1.. Laboratory-scale testing of cracking catalyst. The accurate evaluation of the performance catalytic cracking catalyst and operation conditions is an important issue. As a result of the high throughputs in commercial units small changes in activity and selectivity can influence the probability of an FCC units significantly [1,2]. The deactivation of the catalyst is one of the most critical steps for a realistic evaluation of the catalyst performance [2]. A difference between the reversible and irreversible deactivation should be made. The age distribution and the deposition of metals on the catalyst influence its performance. A newly synthesised catalyst is, therefore, generally artificially deactivated in the laboratory in order to mimic the performance of an equilibrium catalyst from a commercial unit. Most common is the cyclic deactivation in which the catalyst is exposed to cracking and regeneration cycles, using a feed with enhanced metals content [3]. Furthermore, the deposition of coke on the catalyst plays an important role on the activity and selectivity. Especially, the hydrodynamics and mode of operation of laboratory devices are of key importance. Micro-activity testing (MAT) devices are relatively cheap in operation, but the hydrodynamics are not well defined and completely different from commercially applied risers. The principle of operation is that for a given time (time-on-stream) the oil flow is led over a fixed catalyst bed. Major drawbacks are the development of temperature and coke profiles, and channelling in the catalyst bed [4,5]. The selectivities change significantly during the catalyst time-on-stream as a result of the change in intrinsic selectivity of the catalyst by the coke deposition [5]. The residence time (not time-on-stream) is defined as the weight of catalyst divided by the oil injection rate [1]. This implies that, in order to increase the cracking severity, the amount of catalyst and residence time cannot be decoupled, as in commercial units. Fixed fluidised beds provide a better isothermicity and feed dispersion. Moreover, they do not suffer from coke profiles. Backmixing and very long catalyst time-onstream (to obtain fluidisation) are major disadvantages [5,6]. Another option is to pulse the feed over a fixed or fluidised catalyst bed [2,7,8]. An example is a short-contact time residence time test (SCTRT) in which the feed is injected into a fluidised bed of hot catalyst particles followed by fast disengagement of the catalyst and product. The oil time-on-stream is typically one second and simulates adiabatic behaviour at realistic catalyst-to-oil ratios and vapour contact times [2]. Pilot plant riser reactors approach the commercial process more accurately [9,10]. A good example is the Davison Circulating Riser (DCR), which consists of a riser and regenerator can be operated in an adiabatic, isothermal, or pseudo heat balance mode [10,11]. Pilot plant risers give realistic data, but they are more expensive, complex to operate, and require large amounts of catalyst and feed. 22.

(29) Oil Injection Point. Disengager Oven. 5 ml/min. Syringe Pump. Catalyst Storage. Reaction Run. Oil Preheat Oven. Catalyst Preheat Oven. Liquid Recovery. Gasbag. Catalyst Catchpot (Reaction Run). N2 Catalyst Feeder. Reactor Oven. -60 oC -60 oC. Cyclone 1. Catalyst Storage. Pre & Post Run. N2 Stripping Gas. 25 oC. To ventilation. Catalyst Catchpot (Pre & Post Run). Cyclone 2. Figure 1. The microriser. On a smaller scale once-through systems have been developed that simulate riser (or downer) hydrodynamics. These systems use fewer quantities of catalyst and feed and are capable to yield reliable data. For example Corma et al. developed a microdowner unit in order to simulate catalytic cracking reactions in a down-flowing plug-flow regime [Corma]. An oncethrough microriser reactor has been developed at Delft University of Technology, based on a design of Crosfield Chemicals and shown in figure 1[11-14]. The performance of this device has been compared with other testing methods and has proven to generate valuable and realistic data for commercial FCC applications. The development, characteristics, and operation procedures have extensively been described by Helmsing [11] and Den Hollander [14].. 2.. The once-through microriser reactor. 2.1. Hardware. All cracking experiments in this thesis have been performed in the laboratory-scale oncethrough microriser reactor, which is operated in an isothermal plug-flow regime (figure 1). This testing device consists of four different ovens: the reactor oven, the disengager oven, the oil preheat oven, and the catalyst preheat oven. In the reactor oven a looped reactor with an internal diameter of 4.55⋅10-3 m is placed, as represented in figure 2. The length can be varied from 0.2 m to 33.2m. The catalyst is fed together with nitrogen to the preheat oven by means of the catalyst feeder, shown in figure 3.. 23.

(30) Nitrogen. Hopper. Catalyst. Nitrogen. Wheel. Catalyst + Nitrogen. Figure 2. The looped reactor.. Figure 3. The catalyst feeder.. The expanding nitrogen gas in the preheat oven accelerates the catalyst and facilitates the transport to the reaction section. Besides that, the nitrogen acts as a diluting medium that decreases the oil partial pressure in the reactor. With use of a pulse-free constant flow syringe pump (ISCO 500D) the liquid oil feed can be heated (if desired to 110 °C) and transported to the preheat oven. In this preheat oven the feed is partially evaporated at 350 °C and led to the reactor. In the last section before the injection point the feed adapts the reactor temperature and is fully evaporated [14]. The oil is sprayed perpendicular in the down-flowing stream of catalyst and nitrogen. Both feed and catalyst possess the same temperature. The cracking reactions are terminated in the disengager oven, where the catalyst is separated from the gaseous products by means of a cyclone. After the reaction run the valves are switched back and the oil and catalyst flow are turned off (post-run). The oil lines are then purged with nitrogen without affecting the collected sample and stored catalyst. After this cleaning step the stored catalyst is stripped from adsorbed hydrocarbons with nitrogen. An additional step is applied in which the dewars are removed and the condensed product is equilibrated with the ambient conditions under continuous purging with nitrogen. The stored catalyst is removed from the collection vessel and. 24.

(31) transferred to the catalyst catchpots with nitrogen. The product is ready for measurement. In table 1 a summary of the procedure is given. Before the experiments the catalyst is pre-treated with air in a fluidised bed at 600 °C for 180 minutes to combust any coke from the spent equilibrium catalyst. After this pre-treatment step the catalyst is kept at 200 °C. Before placing the catalyst in the feeder it is cooled down to room temperature in vacuum, using a desiccator. 2.2. Flow and temperature profiles. The reactor consists of mountable segments of 2 m in length each. The smallest reactor that can be used in this configuration is 1.2 m; the largest reactor has a length of 33.2 m. In case the reactor of 0.2 m is used the injection point is placed directly upstream of the cyclones in the disengager oven. Due to the choice of the number of loops as a reactor the direction of catalyst- and oil flow changes alternately from up-flow to down-flow. Moreover, the flow through the bends can potentially cause flow profiles that deviate form the desired plug-flow regime. The plug-flow regime can be maintained when the solids in the mixture follow the streamlines of the gas, i.e. when the drag force applied on the catalyst by the gas is sufficient enough to prevent solids backmixing. Helmsing performed several residence time distributions measurements at ambient temperature and defined the regions in which the plug-flow is defined as acceptable. At higher CTO’s higher gas velocities are required to maintain the regime [13,14]. In this thesis it has been verified that the catalyst- and oil feed rates are sufficiently high for this the plugflow regime. Table 1. Microriser operation procedure. Step Reactor purging with nitrogen Glassware purging with nitrogen Pre-run (start-up catalyst and oil flow) Reaction run (steady-state sample collection) Post-run (shutdown catalyst and oil flow) Oil-line purge with nitrogen Catalyst stripping with nitrogen Equilibration with ambient conditions under nitrogen purge. Time [min] 10 10 5 10-15 5 10 20 20. 25.

(32) The initial contact of the feed with the catalyst in the microriser differs from commercial units. In an industrial unit the heat of the catalyst from the regenerator is partially used for the evaporation of the ‘colder’ feed. The catalyst experiences a so-called ‘thermal shock’ [5]. Moreover, the mixing of the catalyst particles with the feed takes place under turbulent conditions and, hence, concentration- and temperature profiles can be expected. The microriser operates as an isothermal cracking device. All the heat for the reactions is provided by the heating elements. The distance that the catalyst and oil travel from the catalyst preheat oven and oil preheat oven, respectively to the injection point is long enough to adapt the reaction temperature and no ‘thermal shock’ is experienced [14].. 3.. Collection and standard analysis of products. 3.1. Liquid product. The three liquid fractions from the glassware are collected and mixed together. The glassware is then rinsed with carbon disulfide (CS2) and is subsequently mixed with collected liquid fractions. Typically, for a 10 minute run 100 ml CS2 is used. Samples were taken and the hydrocarbon composition is analysed with a modified simulated distillation ASTM D2887 method (SimDis). The equipment used is a Hewlett Packard type 5890 series II chromatograph with a HP-1 column (7.5 m x 0.53 mm, film thickness 2.65 µm), using helium as carrier gas and an oven heating program from 35-350 °C at 14 °C/min.. (a) Liquid products LCO. HCO. Cracking product. Signal [a.u.]. Signal [a.u.]. Gln. (b) Gaseous products. Feed. Retention time [a.u.]. (a) Liquid product distribution. C1 -C2. C4. (b) Gas product distribution. Figure 4. Liquid and gas product distributions. 26. C3. C5 -C6.

(33) To this column a flame ionisation detector is connected to determine the hydrocarbon fractions. The injector block is heated during the run from 100-350 °C at 30 °C/min. A boiling point calibration and a reference oil sample are used to calibrate the system. The boiling range for the gasoline fraction is set at C5-215 °C, for light cycle oil (LCO) at 215-325 °C and for high cycle oil (HCO) above 325 °C, respectively. Typical simdis spectra for the feed and cracking product are shown in figure 4a. 3.2. Gaseous product. A Delsi-Nermag type DN200 gas chromatograph is used to analyse the non-condensable gas products (C1-C5) from the Tedlar bag. A 13X molecular sieve packed column (3m x 2mm) with argon as carrier gas is operated isothermally at 60 °C and connected to a thermal conductivity detector (TCD) to detect the hydrogen, methane, and nitrogen content. An OV1 type column (50 m x 320 µm capillary, film thickness 5 µm), with helium as carrier gas, was operated by means of a temperature programme, running from 60 to 195 °C at 6.5 °C/min. To this column a flame ionisation detector (FID) is connected to determine the fraction of hydrocarbon products. The response factors are calculated by calibration of both detectors with a gas cylinder of certified hydrocarbons in the C1-C5 range and nitrogen as a balance gas. An example of the gas distribution is shown in figure 4b (obtained with FID detector). 3.3. Solid product. The coke on the spent catalyst is determined with a LECO C-400 analyser. This device combusts the coke into CO and CO2, and both are detected by an infrared detector. From the known sample weight and a calibration sample the coke content on the catalyst can be calculated. In table 2 the distribution of all products is given.. Table 2. Cutpoints for the hydrocarbon lumping. Product. Composition/Boiling range. Dry gas LPG Gas * Gasoline. H2, C1-C2 C3-C4 H2, C1-C4 C5-215 °C. Light Cycle Oil (LCO). 215-325 °C. Heavy Cycle Oil (HCO). 325+ °C. Coke. Solid carbonaceous species. * Gas is the sum of dry gas and LPG. 27.

(34) 3.4. Calculations. The mass recovery of the experiments is determined by weighing the mass of the liquid and solid products (coke on the catalyst). The mass of the gasbag is calculated from the known amount of nitrogen introduced into the system. A mass recovery of at least 90 % is a prerequisite. The conversion of the feed is based on the amount of HCO converted:. X=. yhco , feed − yhco yhco , feed. ⋅100 %. (1). The residence time is calculated on basis of the nitrogen flow, catalyst flow, and hydrocarbon volumetric flow under reaction conditions (Treaction and Preaction) [14,15]:. τ=. 0.25 ⋅ π ⋅ d 2 ⋅ L φv , N2 + φv , HC + φv ,Catalyst. (2). The volumetric flow rates of nitrogen and the hydrocarbons of the gas fraction are calculated via the analysis with the GC. The hydrocarbon volumetric flow of HCO, LCO, and gasoline, based on outlet conditions, are calculated by:. φv , HC =. φ R ⋅ TR ⋅ ∑ m ,i pR M w ,i. (3). The molecular weight of HCO, LCO, and gasoline are set at 350, 200, and 100 g/mol, respectively [15].. 28.

(35) 4.. Symbols and Abbreviations. CTO. catalyst-to-oil ratio. [gcat⋅gfeed-1]. d HCO L LCO. reactor internal diameter heavy cycle oil reactor length light cycle oil. [m] [-] [m] [-]. Mw,i. molecular weight. [g⋅mol-1]. pR. reactor outlet pressure. [Pa]. R. gas constant. [J⋅mol-1⋅K-1]. T X yi. reactor temperature HCO conversion product mass fraction. [K] [wt%] [-]. Φm,i. mass flow rate. [kg⋅s-1]. Φv,i. volumetric flow rate. [m3⋅s-1]. τ. residence time. [s]. 29.

(36) 5.. References. [1] [2] [3]. C.P. Kelkar, M. Xu, and R.J. Madon, Ind. Eng. Chem. Res. 42, 426-433 (2003) P. Imhof, M. Baas, and J.A. Gonzalez, Catal. Rev. 46, 151-161 (2004) D.R. Rainer, E. Rautiainen, B. Nelissen, P. Imhof, and C. Vadovic, Stud. Surf. Sci. Catal. 149, 165-176 (2004) A. Corma, C. Martínez, F.V. Melo, L. Sauvanaud, and J.Y. Carriat, Appl. Catal. A. 232, 247-263 (2002) D. Wallenstein, A. Haas, and R.H. Harding, Appl. Catal. A. 203, 23-36 (2000) R.C. Vieira, J.C. Pinto, E.C. Biscaia Jr., C.M.L.A. Baptista, and H.S. Cerqueira, Ind. Eng. Chem. Res. 43, 6027-6034 (2004) J.A. Atias and H. de Lasa, Chem. Eng. Sci. 59, 5663-5669 (2004) K. Lipiänen, P. Hagelberg, J. Aittamaa, I. Eilos, J. Hiltunen, V.M. Niemi, and A.O.I. Krause, Appl. Catal. A. 183, 411-421 (1999) G.M. Bollas, I.A. Vassalos, A.A. Lappas, D.K. Iatridis, and G.K. Tsioni, Ind. Eng. Chem. Res. 43, 3270-3281 (2004) G.W. Young, Stud. Surf. Sci. Catal. 76, 257-292 (1993) M.P. Helmsing, “FCC catalyst testing in a novel laboratory riser reactor”, Ph.D. Thesis, Delft University of Technology, Delft (1996) D.J. Rawlence and K. Gosling, Appl. Catal. 43, 213-237 (1988) M.P. Helmsing, M. Makkee, and J.A. Moulijn, Chem. Eng. Sci. 51, 3039-3044 (1996) M.A. den Hollander, “Catalytic cracking in a microriser, the different time scales of cracking and coke deposition”, Ph.D. Thesis, Delft University of Technology, Delft (2000) M.A. Den Hollander, M. Makkee and J.A. Moulijn, Appl. Catal. A. 187, 3-12 (1999). [4] [5] [6] [7] [8] [9] [10] [11] [12] [13] [14]. [15]. 30.

(37) 3 Cracking behaviour of organic sulfur compounds under realistic FCC conditions Abstract A qualitative study under realistic FCC conditions on the distribution of aromatic sulfur in the catalytic cracking product of an extra heavy gas oil (EHGO) has been performed. Fifteen to twenty percent of the feed is uncrackable, corresponding to the amount of polycyclic aromatics in the feed. Gasoline is not overcracked to gas. In contrast to the selectivity to LCO and gas, the gasoline selectivity is independent of process parameters. Initially, the amount of gasoline sulfur increases with residence time due to the formation of aliphatic sulfur compounds, which are subsequently cracked to hydrocarbons and H2S. The gasoline sulfur concentration decreases as function of conversion due to the dilution of sulfur compounds by the additional formed sulfur-free gasoline hydrocarbons. Stable LCO sulfur species are predominantly formed by dealkylation of alkyl-benzothiophenes from the HCO fraction. The LCO sulfur concentration is temperature dependent. Aromatic cores without side chains, such as thiophene, benzothiophene, and dibenzothiophene are uncrackable under the applied conditions.. 31.

(38) 1.. Introduction. As a result of the global environmental legislation refiners are faced with the demand of ongoing reduction of sulfur- and aromatic containing components in fuels. In 2005, the limits for maximum sulfur concentration in gasoline and diesel in Europe is 50 ppm, and will be lowered in the forthcoming years to 10 ppm or even lower [1-6]. In automotive engines the sulfur containing compounds are combusted and emitted as SOx [6]. Besides the direct harmful effect of SOx on the environment, it acts as a poison for the noble metal catalysts of exhaust gas after-treatment devices [7-9]. At the request of the automotive industry the future legislative requirements will, therefore, be sulfur free fuels (<10 ppm sulfur). This will not only lead to the direct reduction of SOx- and soot emission, but also will enable particulate-, NOx- and hydrocarbon abatement systems to operate much more efficiently [1,7-10]. The Fluid Catalytic Cracking (FCC) process is the main source for sulfur in gasoline and diesel and, therefore, it could play an important role in the reduction of sulfur concentrations. The FCC unit produces about 40 v% of the total gasoline pool and can contribute to more than 90 % of the sulfur containing components in the overall gasoline product [2,7-15]. Besides that, FCC product light cycle oil (LCO) is an important diesel blend, which may contribute up to 40 v% of the diesel pool (together with thermal cracking and hydrocracking), and is responsible for 75 % of the total sulfur amount [16]. In crude oil a broad spectrum of sulfur-containing compounds are present. Typically, the FCC feed consists of 3 % of sulfur-containing compounds [3]. Most abundant (up to 60 %) are thiophenes, benzothiophenes, and dibenzothiophenes to which alkyl-branches are connected as shown in figure 1 [7,8]. The aromatic character of the thiophene ring makes that these compounds hardly crack under industrial FCC conditions [3,7,8,17]. Only a minor part of the feedstock consists of mercaptanic sulfur species. For the reduction of the sulfur concentration in fuels three different options are open: 1) prevention of the formation of the gasoline- and diesel range sulfur components; 2) pretreatment of the FCC feed; and 3) removal of the sulfur components from the gasoline and LCO fraction [7]. Prevention of the formation of gasoline- and diesel range sulfur components may be economically more attractive than removing sulfur components from the FCC feed or the product fractions. Although hydroprocessing of the FCC feed is a very effective way of sulfur reduction, the high hydrogen pressure and temperature makes this a costly operation [8,9]. Desulfurisation of the product streams may result in a more efficient use of hydrogen, but the side effect is that it can strongly affect fuel composition and quality, especially the gasoline octane number [7].. 32.

(39) S. S R. R. (a). S. R. R. (b). (c). Figure 1. Aromatic sulfur molecules. (a) alkyl-dibenzothiophene, (b) alkyl-benzothiophene, and (c) alkyl-thiophene.. Monitoring the formation of the sulfur components in the fuels range in the FCC unit may result in better understanding and subsequent control over the sulfur concentrations in the fuels. It has been reported that hydride transfer (or hydrogen transfer) will play an important role in the conversion of aromatic sulfur compounds [8,12,14,18,19]. Alkemade et al. [14] have stated that catalysts with low hydride transfer activity yield more sulfur in gasoline than catalysts with high hydride transfer activity. Corma et al. [12,19] have reported on the mechanisms of sulfur compound formation and decomposition, using a micro-activity (MAT) set-up. In this fixed bed reactor at 510 °C and 30 s catalyst time-on-stream, they found that hydride transfer plays an important role in the cracking of sulfur species. Furthermore, initially formed aliphatic sulfur compounds were cracked relatively easy to olefins and H2S, while benzothiophenes are hardly cracked [19]. The objective is to investigate in a qualitative way whether under realistic FCC conditions aromatic sulfur species can be overcracked to H2S or forced to accumulate in the less valuable HCO fraction by variation the process parameters catalyst-to-oil ratio (CTO), residence time (τ), and reaction temperature (T).. 2.. Experimental. The experiments have been performed with an industrial extra heavy gas oil (EHGO), which has been used to evaluate the cracking behaviour of 'aromatic' sulfur with regard to the aromatic properties of the feed. The basic properties of the feed are given in table 1. The used catalyst is a commercial equilibrium catalyst with an average particle size of 78 µm. The catalytic cracking experiments have been performed in the microriser reactor. The catalyst to oil ratio (CTO) has been varied between 2 - 8 gcatalyst⋅goil-1, keeping the oil feed rate constant. The residence time has been varied by changing the reactor length, 1.2 - 21.2 m which roughly corresponds a residence time of 0.3 - 5 s, respectively. The reactor temperatures have been set at 525, 555, and 585 °C.. 33.

(40) Table 1. Properties of the extra heavy gas oil feed. Density at 15 °C. [kg/l]. 0.8784. 2. Viscosity at 100 °C. [m /s]. 2.75⋅10-6. Average molecular weight Sulfur content Nitrogen content. [g/mol] [wt%] [wt%]. 299 0.38 0.028. Carbon content Hydrogen content. [wt%] [wt%]. 86.92 12.72. Distillation curve: [wt%] 0 251 [°C]. 10 309. 20 325. 40 348. 60 369. 80 393. 90 411. 100 524. A five-lump model by Corella and Francés [20], adapted by Den Hollander et al. [21,22], has been used to evaluate the results and is shown in figure 2. The conversion of the feed is based on the amount of HCO converted:. X=. yhco , feed − yhco yhco , feed. ⋅100 %. (1). Aside the standard analysis methods described in chapter 2 a Sievers 355 chemiluminescence detector has been used to determine the sulfur content in the pure liquid samples. This detector is coupled with a Hewlett Packard type 5890 series II gas chromatograph with a BPX05 (60 x 0.22 mm, coated with a 0.25 µm thick fused silica bonded phase). The oven is heated from 150-300 °C at 7.5 °C/min with a dwell time of 20 minutes. An external standard sample (0.01 wt% thiophene in toluene) has been used to calculate the response factor for the chemiluminescence detector. The SimDis and Sievers analyses have been used to determine the feed composition, given in table 2.. Coke R7. R8 R1. HCO R2 Gasoline. R3. LCO R5. R6. R4 Gas. Figure 2. Five-lump model for catalytic cracking [20-22].. 34.

(41) Table 2. Overall- and sulfur distribution of the extra heavy gas oil feed as determined by SimDis and Sievers analysis. Boiling-range. Amount of sulfur [wt%]. Sulfur concentration. [°C]. Product fraction [wt%]. Gasoline. 25-215. 1.1 (± 0.1). 0.003 (± 0.001). 3234 (± 537). LCO. 215-325. 18.8 (± 1.4). 0.022 (± 0.003). 1178 (± 91). HCO. > 325. 80.2 (± 1.5). 0.343 (± 0.011). 4281 (± 63). 3.. [ppm]. Results. In figure 3a the HCO fraction versus residence time is displayed for CTO 2. Through the data points lines are drawn to guide the eye. During the first two seconds most of the HCO has been converted after which the curves exhibit an asymptotic behaviour. The catalytic cracking rate increases with temperature. Approximately 15-20 % of the feed is uncrackable under the investigated conditions. The asymptotic behaviour and temperature dependency has also been observed for the production of gasoline in figure 3b. The gasoline fraction shows a strong increase that is most profound during the first two seconds and then starts to saturate to approximately 50 wt%. No overcracking of gasoline has been observed. In figure 4 the influence of the CTO is presented. The presence of more catalyst enhances the conversion rate and also in this case no overcracking of gasoline has been observed.. 90. 60. (a) HCO (325+ o C). 80. 585 555 525. 50 Gasoline fraction [wt%]. 70 HCO fraction [wt%]. (b) Gasoline (C 5 -215 o C). 60 50 40 30 20. 525 555 585. 10. 40 30 20 10. 0. 0 0. 1. 2 3 4 Residence time [s]. 5. 6. 0. 1. 2 3 4 Residence time [s]. 5. 6. Figure 3. HCO- and gasoline fraction versus residence time at different temperatures and CTO 2. (a) HCO and (b) gasoline. C = feed, ¡= 525 °C, = 555 °C, and z = 585 °C.. 35.

(42) 90. 60. (a) HCO (325 o C). 80. CTO 4. 50 Gasoline fraction [wt%]. 70 HCO fraction [wt%]. (b) Gasoline (C 5 -215 o C). 60 50 40 30 20. CTO 4. 10. CTO 2. CTO 2 40 30 20 10. 0. 0 0. 1. 2 3 4 Residence time [s]. 5. 6. 0. 1. 2 3 4 Residence time [s]. 5. 6. Figure 4. HCO- and gasoline fraction versus residence time at 525 °C for CTO 2 and CTO 4. (a) HCO and (b) gasoline. C = feed, ¡= CTO 2 and ▲= CTO 4.. In figure 5 the product fractions versus HCO conversion are represented. These graphs show that the gasoline fraction is a linear function of conversion. The LCO- and gas fraction are temperature dependent. The LCO fraction displays a maximum at a conversion of 40-50 %, which lies higher at lower temperatures. For the gas fraction (H2, C1-C4) the opposite is the case, i.e. with increasing temperature the gas fraction increases. The offset of the gas fractions from the origin can be explained by the fact that during the injection of the feed small gas molecules are formed as a result of thermal cracking and the stripping of coked catalyst. During the injection of the feed perpendicular into the catalyst flow the hydrodynamics are partially disturbed. For the studied temperatures the intersection with the y-axis is the same (~ 2 wt% gas), i.e. the effect of thermal cracking is low. The coke fraction is a scatter plot from which no clear trend can be observed. In figure 6 the hydrocarbon distribution, as obtained from the simulated distillation, is shown. With increasing conversion level the HCO fraction decreases and a significant increase of low-boiling hydrocarbons in the gasoline range is observed. In the LCO fraction there is a shift from high-boiling hydrocarbons to low-boiling hydrocarbons. At all conversions the fingerprint of the spectrum is similar, i.e. only the intensities of the peaks change. In figure 7 the amount of sulfur versus residence is shown at CTO 2. Figure 7a shows the fast cracking of sulfur compounds from the HCO range during the first two seconds of residence time. After that the amount of HCO sulfur remains constant at approximately 0.17 wt%.. 36.

(43) 30. 60. (a) LCO (215-325 o C). 50 525. 20. 555 585. 15 10. Gasoline fraction [wt%]. LCO fraction [wt%]. 25. 5. 40 30 20 10. 0. 0 0. 20. 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. 0. 5. (c) Gas (H 2 ,C 1 -C 4 ) 585. 16. 555. 14 12. 525. 10 8 6. 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. (d) Coke fraction vs. conversion. 4 Coke fraction [wt%]. 18. Gas fraction [wt%]. (b) Gasoline (C 5 -215 o C). 3. 2. 1. 4 2 0. 0 0. 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. 0. 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. Figure 5. Product fractions versus conversion. (a) LCO, (b) gasoline, (c) gas, and (d) coke. C = feed, ¡= 525 °C, = 555 °C, and z = 585 °C.. LCO sulfur is formed rapidly during the first second of reaction and then saturates at approximately 0.13 wt%. For the amount of gasoline sulfur a maximum is observed. The total amount of sulfur in the liquid recovery decreases with larger residence time, at least partially due to formation of gaseous sulfur species. In figure 8 the amount of sulfur species versus conversion is given. For all three temperatures the amount of HCO sulfur decreases with increasing conversion. The amount of LCO sulfur increases and the slope approaches zero for high conversion. The amount of gasoline sulfur goes through a maximum at 40-50 % HCO conversion.. 37.

(44) (a) Hydrocarbon distribution of the feed LCO. HCO. Gln. LCO. HCO. Signal [a.u.]. Signal [a.u.]. Gln. (b) Distribution at X=51.3 %. Retention time [a.u.]. (c) Distribution at X=65.6 %. (d) Distribution at X=78.5 %. HCO. Gln. LCO. HCO. Signal [a.u.]. LCO. Signal [a.u.]. Gln. Retention time [a.u.]. Retention time [a.u.]. Retention time [a.u.]. Figure 6. Hydrocarbon distribution at different HCO conversion levels. (a) feed, (b) X= 51.3 %, (b) X= 65.6 %, and (c) X= 78.5 % (all at 585 °C).. The gasoline- and LCO sulfur data from figure 8 have been combined in figure 9. From this figure it can be concluded that no significant temperature dependency for the amount of gasoline- and LCO sulfur can be found. In figure 10a-c the sulfur distribution is represented at 555 °C and CTO 2. In these spectra thiophene, benzothiophene, and dibenzothiophene (DBT) have been identified. It is clear that most of the sulfur in the feed lies in the HCO range and only a minor part in the LCO range. The spectrum consists of sulfur species with a variety of boiling points that lie close together.. 38.

(45) 0.40. 0.30 0.25 0.20 0.15 0.10 0.05. 0.12 0.10 0.08 0.06 0.04 0.02. 0.00. 0.00 0. 1. 2 3 4 Residence time [s]. 5. 6. 0. 0.035. 0.45. (c) Gasoline sulfur amount vs. residence time 0.030 0.025 0.020 0.015 0.010 0.005. 1. 2 3 4 Residence time [s]. 5. 6. (d) Total sulfur amount vs. residence time. 0.40 Total sulfur amount [wt%]. Gasoline sulfur amount [wt%]. (b) LCO sulfur amount vs. residence time. 0.14 LCO sulfur amount [wt%]. 0.35 HCO sulfur amount [wt%]. 0.16. (a) HCO sulfur amount vs. residence time. 0.35 0.30 0.25 0.20 0.15 0.10 0.05. 0.000. 0.00 0. 1. 2 3 4 Residence time [s]. 5. 6. 0. 1. 2 3 4 Residence time [s]. 5. 6. Figure 7. Absolute amount of sulfur versus residence time at CTO 2. (a) HCO range sulfur, (b) LCO range sulfur, (c) gasoline range sulfur, and (d) total amount of sulfur (HCO + LCO + gasoline). C = feed, ¡= 525 °C, = 555 °C, and z = 585 °C.. The resolution of the Sievers analysis method is not high enough to differentiate between all the different sulfur compounds. At X= 52.2 % already a large amount of HCO sulfur has been converted to LCO- and gasoline sulfur. At X= 82.0 % even more HCO sulfur has been converted. It should be noted that the spectrum itself does not change for higher conversion, i.e. no new sulfur components are found. All the sulfur species present at X= 82.0 % are already present at X= 52.2 %. In figure 10d the sulfur distribution is shown at 555 °C and CTO 4. This diagram shows that the spectrum itself is similar to the spectrum at CTO 2. In figure 11 the sulfur distribution is shown at 525 °C and 585 °C, both at CTO 2.. 39.

(46) 0.40 0.35. 0.25 0.20 0.15 0.10. 0.30 0.25 0.20 0.15 0.10. 0.00. 0.00 0. Sulfur amount [wt%]. (b) Sulfur amount at 555 o C. 0.05. 0.05. 0.35. 0.35 Sulfur amount [wt%]. Sulfur amount [wt%]. 0.30. 0.40. 0.40. (a) Sulfur amount at 525 o C. 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. 0 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. (c) Sulfur amount at 585 o C. 0.30 0.25 0.20 0.15 0.10 0.05 0.00 0 10 20 30 40 50 60 70 80 90 100 HCO Conversion [wt%]. Figure 8. Absolute amount of sulfur versus conversion at (a) 525 °C, (b) 555 °C, and (c) 585 °C. ¡= HCO range sulfur, = LCO range sulfur, ▲= gasoline range sulfur, ±= HCO + LCO + gasoline. Open symbols represent the amount of sulfur in the feed, determined by Sievers analysis, except for {= amount of sulfur determined by supplier.. All the sulfur spectra are similar from qualitative point of view, regardless the different applied temperatures, CTO’s, and residence times. In figure 12 the sulfur concentrations versus conversion are shown. For HCO the sulfur concentration is constant up to approximately 50 % conversion. At a higher conversion level a sharp increase of HCO sulfur concentration has been found. For LCO sulfur a temperature dependency has been observed. With increasing temperature the concentration increases.. 40.