Catalyst deactivation particles

Pore volume is an indication of the quantity of voids in the catalyst particles and can be a clue in detecting the type of catalyst deactivation that takes place in a commercial unit. Hydrothermal deactivation has very little effect on pore volume, whereas thermal deactivation decreases pore volume. 

Chiral lactones were also obtained by cyclocarbonylation of chiral acetylenic alcohols with Pd and thiourea (H2NCSNH2) (Scheme 32). No loss in chirality was observed, but large amounts of Pd and thiourea were used (10 mol %) since the catalyst deactivates by forming metal particles. The catalytic precursor (Pdl2 > PdCl2) and the ratio of thiourea to Pd were very important, thiourea being necessary for this reaction. The active species was supposed to be [Pd(thiourea)3l]I, which forms in situ from [Pd(thiourea)4]l2 and [Pd(thiourea)2]l2. It had to be a partially dissociated species since [Pd(thiourea)4](Bp4)2 was inactive [121]. 

Reaction, diffusion, and catalyst deactivation in a porous catalyst layer are considered. A general model for mass transfer and reaction in a porous particle with an arbitrary geometry can be written as follows ... 

Figure 8.8 Internal mass transfer resistance and catalyst deactivation concentration profiles inside a catalyst particle-lactose hydrogenation to lactitol and by-products (sponge Ni).

The catalyst deactivates, but after four runs the conversion is still significantly higher (> 99% after 2 h) as compared with the uncatalyzed reaction. Moreover, the Z-selectivity in all four runs is higher than 80%, whereas in the uncatalyzed reaction, it is typically only 30% (Z). The fact that the solid powder can be used several times furthermore supports the fact that the reaction mechanism is heterogeneous. The reason for the deactivation is unknown. A disadvantage of the nanoparticles is the difficulty of separation. Thus, in some cases the particles form col-... 

The activity and stability of catalysts for methane-carbon dioxide reforming depend subtly upon the support and the active metal. Methane decomposes to carbon and hydrogen, forming carbon on the oxide support and the metal. Carbon on the metal is reactive and can be oxidized to CO by oxygen from dissociatively adsorbed COj. For noble metals this reaction is fast, leading to low coke accumulation on the metal particles The rate of carbon formation on the support is proportional to the concentration of Lewis acid sites. This carbon is non reactive and may cover the Pt particles causing catalyst deactivation. Hence, the combination of Pt with a support low in acid sites, such as ZrO, is well suited for long term stable operation. For non-noble metals such as Ni, the rate of CH4 dissociation exceeds the rate of oxidation drastically and carbon forms rapidly on the metal in the form of filaments. The rate of carbon filament formation is proportional to the particle size of Ni Below a critical Ni particle size (d<2 nm), formation of carbon slowed down dramatically Well dispersed Ni supported on ZrO is thus a viable alternative to the noble metal based materials. 

The selectivity in a system of parallel reactions does not depend much on the catalyst size if effective diffusivities of reactants, intermediates, and products are similar. The same applies to consecutive reactions with the product desired being the final product in the series. In contrast with this, for consecutive reactions in which the intermediate is the desired product, the selectivity much depends on the catalyst size. This was proven by Edvinsson and Cybulski (1994, 1995) for. selective hydrogenations and also by Colen et al. (1988) for the hydrogenation of unsaturated fats. Diffusion limitations can also affect catalyst deactivation. Poisoning by deposition of impurities in the feed is usually slower for larger particles. However, if carbonaceous depositions are formed on the catalyst internal surface, ageing might not depend very much on the catalyst size. 

Batchwise operating three-phase reactors are frequently used in the production of fine and specialty chemicals, such as ingredients in drags, perfumes and alimentary products. Large-scale chemical industry, on the other hand, is often used with continuous reactors. As we developed a parallel screening system for catalytic three-phase processes, the first decision concerned the operation mode batchwise or continuous. We decided for a continuous reactor system. Batchwise operated parallel sluny reactors are conunercially available, but it is in many cases difficult to reveal catalyst deactivation from batch experiments. In addition, investigation of the effect of catalyst particle size on the overall activity and product distribution is easier in a continuous device. 

In a fixed-bed catalytic reactor for a fluid-solid reaction, the solid catalyst is present as a bed of relatively small individual particles, randomly oriented and fixed in position. The fluid moves by convective flow through the spaces between the particles. There may also be diffusive flow or transport within the particles, as described in Chapter 8. The relevant kinetics of such reactions are treated in Section 8.5. The fluid may be either a gas or liquid, but we concentrate primarily on catalyzed gas-phase reactions, more common in this situation. We also focus on steady-state operation, thus ignoring any implications of catalyst deactivation with time (Section 8.6). The importance of fixed-bed catalytic reactors can be appreciated from their use in the manufacture of such large-tonnage products as sulfuric acid, ammonia, and methanol (see Figures 1.4,11.5, and 11.6, respectively). 

Industrially relevant consecutive-competitive reaction schemes on metal catalysts were considered hydrogenation of citral, xylose and lactose. The first case study is relevant for perfumery industry, while the latter ones are used for the production of sweeteners. The catalysts deactivate during the process. The yields of the desired products are steered by mass transfer conditions and the concentration fronts move inside the particles due to catalyst deactivation. The reaction-deactivation-diffusion model was solved and the model was used to predict the behaviours of semi-batch reactors. Depending on the hydrogen concentration level on the catalyst surface, the product distribution can be steered towards isomerization or hydrogenation products. The tool developed in this work can be used for simulation and optimization of stirred tanks in laboratory and industrial scale. 

Some simulation results for trilobic particles (citral hydrogenation) are provided by Fig. 2. As the figure reveals, the process is heavily diffusion-limited, not only by hydrogen diffusion but also that of the organic educts and products. The effectiviness factor is typically within the range 0.03-1. In case of lower stirrer rates, the role of external diffusion limitation becomes more profound. Furthermore, the quasi-stationary concentration fronts move inside the catalyst pellet, as the catalyst deactivation proceeds. 

The importance of catalyst stability is often underestimated not only in academia but also in many sectors of industry, notably in the fine chemicals industry, where high selectivities are the main objective (1). Catalyst deactivation is inevitable, but it can be retarded and some of its consequences avoided (2). Deactivation itself is a complex phenomenon. For instance, active sites might be poisoned by feed impurities, reactants, intermediates and products (3). Other causes of catalyst deactivation are particle sintering, metal and support leaching, attrition and deposition of inactive materials on the catalyst surface (4). Catalyst poisons are usually substances, whose interaction with the active surface sites is very strong and irreversible, whereas inhibitors generally weakly and reversibly adsorb on the catalyst surface. Selective poisons are sometimes used intentionally to adjust the selectivity of a particular reaction (2). 

Coke burning during regeneration. Co F-T catalysts deactivate due to coke formation blocking the active sites. This coke can be burned off by an oxidative treatment. The addition of promoter elements may decrease the temperature of this oxidative treatment, preventing the possible clustering of supported cobalt particles. 

Fig. 12. Effect of self-poisoning (i hr at 450 C, by the reaction mixture) on hexane/H2 reactions. AT is the temperature increase necessary to achieve the same overall conversion after poisoning over that before poisoning. AT is plotted as a function of the average particle size of various Pt/SiO2 catalysts. From V. Ponec et al, in Catalyst Deactivation, p. 93, Elsevier, Amsterdam (1980).

For example, the most noteworthy disadvantage of catalytic wet oxidation is the severe catalyst deactivation (Larachi el al., 1999). Hamoudi el al. (1998, 1999) systematically studied the deactivation of Mn02/Ce02 catalyst during wet catalytic oxidation of phenol and the catalyst-surface modifications. It was observed that deactivation was induced mainly by the formation of carbonaceous deposits on the catalyst surface. Ohta et al. (1980) reported that the size of the catalyst particles affected the stabilization of catalytic activity. For granular particles of supported copper oxide, the catalytic activity was decreased after each inn, even after six successive experiments. In contrast, for larger particles the catalytic activity was stabilized after the first three runs. 

The present authors proposed ferromagnetic supports, which can withstand liquefaction environment and conditions by their inherent nature or by protection with a carbon coating (141). Carbon has been recognized as an excellent support for FeS and Ni-MoS as described in the previous section (50). Very fine particles of ferrite are available. Coating with carbon can be performed through precipitation of polymers or pitches by the aid of suspension agents followed by carbonization (Ml). Catalyst deactivation by mineral and carbon deposition should be avoided for this approach to be feasible. 

The route of catalyst deactivation via a cyclic metal impregnation and deactivation method has produced significant improvements in approaching realistic vanadium and nickel profiles over the catalyst particles. From electron microprobe analyses of Ni and V loaded catalyst it has been established that after pore volume saturation, Ni and V are rather homogeneously distributed over the catalyst. 

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