Catalysts deactivation rates

Several processes based on non-precious metal also exist. Because of high catalyst deactivation rates with these catalyst systems, they all require some form of continuous regeneration. The Fluid Hydroforming process uses fluid solids techniques to move catalyst between reactor and regenerator TCR and Hyperforming use some form of a moving bed system. 

The average catalyst deactivation rate over the entire experiment was 0.0291%/mscf/lb. The rate of deactivation during the initial 462 hrs of operation at a fresh feed space velocity of about 2090/hr (216 scfh) was very low, 0.0017%/mscf/lb from 500 hrs to 841 hrs with 2990/hr space velocity, the deactivation rate increased to 0.040% /mscf/lb. Catalyst deactivation rates were calculated (Table IX) for various operating periods and fresh feed space velocities. 

Table IX. Catalyst Deactivation Rates during Experiment HGR-14...

One of the major disadvantages of fixed bed operation is that catalyst regeneration or replacement is relatively difficult to accomplish. If the catalyst deactivation rate is sufficiently... 

Smith and co-workers investigated the effect of metal support interaction on the CF formation on a series of Co-silica catalysts.The metal support interaction was manipulated by addition of either BaO, La203 or Zr02 to silica. The rate of catalyst deactivation was found to increase with the increase in the metal support interaction. Competition between CF formation and encapsulating carbon formation controlled the catalyst deactivation rate. In case of the catalysts with high metal support interaction, the encapsulating carbon formation was dominant and hence led to a rapid deactivation of the catalyst. 

Catalysts in coal liquefaction are used in moving-bed, ebulating-bed, and fixed-bed processes. Disposable iron catalysts must be used in moving beds. More expensive Co-Mo and Ni-Mo catalysts are used in either ebulating or fixed beds, and catalyst deactivation rates and ultimate lifetime are of concern (80, 81). In ebulating beds, a small portion of fresh catalyst is continuously fed to balance the catalyst being purged. 

The severity of the conditions can also influence the catalyst deactivation rate. Higher severity or deeper desulfurization can be obtained by operation at higher temperatures for more catalyst activity. As shown in Fig. 54, however, this higher activity is at the expense of useful catalyst life. At higher temperature the metals distribution parameter is lower and coke formation is more rapid because of increased catalyst activity. 

Higher pressure operation in the LC-Finer tends toward a decrease in the catalyst deactivation rate for conversion. This is believed to be due to a lower coking rate at a higher hydrogen partial pressure. 

Figure 4 Impact of catalyst deactivation rate on predicted MAT and FFB conversions.

This decrease in residence time led to a smaller catalyst inventory in the reactor-regenerator system, and consequently much lower catalyst deactivation rates due to hydrothermal effects. A lower catalyst inventory has also the additional operational advantage that the FCCU can respond more quickly to changes in catalyst management, allowing the refiner to restore the desired catalyst activity level more promptly following an upset. 

Generally, an amount of coke on the catalyst increases from the entrance to the exit of the fixed bed reactors in residue hydroprocessing (1, 6, 7). Tamm et al. showed the highest remained catalyst activity at the outlet of the bench-scale fixed-bed reactor after a constant desulfurization operation (1), while Myers et al. found the highest catalyst deactivation rate in the last stage of three-stage pilot-scale expanded-bed reactors after a 60 - 70% vacuum residue conversion operation (7). These results from two typical reactor operations support that the catalyst deactivation in a lower... 

Although a great deal is known about catalyst deactivation in conventional, fixed-bed reactors, several new issues arise when a slurry reactor is used at conditions that are representative of the CGCC application. These include 1) the use of a concentrated. CO-rich feed gas 2) the likely presence of relatively-high concentrations of metal carbonyls, and 3) the need to define very precisely the effect of temperature on catalyst productivity and catalyst deactivation rate. The effect of feed gas composition and poisoning by carbonyl compounds was previously discussed [ref. 3[. This paper is concerned with the effect of temperature. 

Deactivation rates and aged catalyst properties have been investigated as a function of time on stream for iron-based Fischer-Tropsch catalysts in the presence/absence of potassium and/or silicon. There is a synergism in activity maintenance with the addition of both potassium and silicon to an iron catalyst. The addition of silicon appears to stabilize the surface area of the catalyst. Catalysts containing only iron or added silicon with or without potassium consist mainly of iron oxide at the end of the run. However, iron carbides are the dominant phase of the iron catalyst with added potassium alone. Catalyst surface areas increase slightly during synthesis. The bulk phase of the catalyst does not correlate to the catalyst activity. The partial pressure of water in the reactor is lower for potassium-containing catalysts and is not a reliable predictor of catalyst deactivation rate. 

There were two major objectives of this study. Firstly, the effects of the addition of potassium and/or silicon on catalyst deactivation rates and changes in catalyst properties with TOS were investigated. Secondly, the possible causes of catalyst deactivation were examined by following aged catalyst properties and reactor conditions as a function of TOS for each catalyst. The FTS was carried out in a continuous-flow stirred slurr) reactor to ensure uniformity in catalyst aging and reactor conditions throughout the reactor. The aged catalyst properties examined as a ftinction of TOS were total surface areas, carbon deposits and phase transformations. 

Since the n-heptane reforming rate is much faster than catalyst deactivation rate, the reaction system considered, with appropriate assumptions, is represented by the following quasi-steady state mass balance equations ... 

An initial period of stabilization (about two months) was experienced with this catalyst system, when high (8°F/month) catalyst deactivation rate was observed. After this, a period of low catalyst deactivation rate followed (4-5°F/ month WABT). Regarding the reactors exotherms, the same trend as with Catalyst System-B was observed. First main reactor showed about 50% of total delta temperature which drifted to the next reactor at EOR. 

Figure 4 is an example of a long-term bench plant test for a catalyst combination system. Several ten days after the start-of-run, the catalyst system showed stable deactivation. During a stable deactivation period, the catalyst deactivation rate is constant. If the operation mode was a constant product sulfur mode, the temperature-increase-rate of reaction (TIR) was constant and small. Then, after the stable deactivation period, the catalyst system showed a higher deactivation period, in which the TIR became constantly larger than that during the stable deactivation period. The point at which the deactivation rate changes is called a breakpoint. 

The catalyst deactivation rate decreased when CO2 was introduced into the system with a very law p ial pressure. These results indicate that the carbonate species were formed instead of the volatile species. The carbonate species are more stable but less catalytically active. 

A typical EBOne plant flow diagram is shown in Fig. 4. The alkylation reactor is maintained in the liquid phase and uses multiple catalyst beds and ethylene injections to improve the reaction selectivity. Dividing the ethylene into multiple feed streams keeps the alkylation catalyst deactivation rate very low. In some plants using EBZ-500 catalyst, operating lengths of more than 8 yr have been obtained without catalyst regeneration. The ethylene conversion is essentially 100% in the alkylation reactors, and the reactors operate nearly adiabatically. The exothermic heat of reaction is recovered and used within the process to heat internal process streams or to generate steam. 

Silylated catalysts prepared from methoxytrimethylsilane [71,85] contacted with amorphous, disordered mesoporous titanosUicate (2 mol%) followed by deposition of Au via DP resulted in an increase in PO production rate from 52 to 67 gpo kg f h over the unsilylated catalyst at 150 °C. There was no significant effect on H2 efficiency (an increase from 33.3% to 35.3%). Catalyst deactivation rate appears to have been decreased slightly, 58% retention of activity over the first 4 h relative to 44% for the unsilylated catalyst. SUylation in combination with Ba promotion and the high Ti content of the disordered mesoporous materials has resulted in some of the most active catalysts thus reported with a PO rate of 92 gpo kg j h at 150 °C. 

In response to this limitation, several processes have been developed that suppress catalyst deactivation through the use of slurry reactors, catalytic distillation, and an FCC-style fluidized bed with constant catalyst regeneration (6-8). These reactors achieve high I/O ratios and lower catalyst deactivation rates, but at a cost. Fluidized beds, for example, require specialized, attrition-resistant catalyst and extensive filtration equipment. 

We have shown that additive coke (Cat]j) has much less impact on catalyst activity than catalytic coke (Ccal) at the same coke-on-catalyst level, but the initial catalyst deactivation rate during ARO cracking is greaier than ihat of VGO cracking because of the fast deposition of additional coke on the caialyst surface. The general correlations developed in this paper can be conveniently u sod in the modeling of catalytic cracking reaction kinetics. 

Finally, The dehydrogenation of butanediols to y-butyrolactone is an important commercial reaction that was developed by BASF and named the Reppe process. The most probable reaction mechanism via the y-hydroxybutyraldehyde intermediate clearly shows that the reaction proceeds via two separate alcohol dehydrogenation steps with a rearrangement step taking place in-between (Table 1, Scheme 12) [49]. The reaction is usually performed in the gas phase with hydrogen as carrier gas, to reduce catalyst deactivation, which is a characteristic problem. Thus, extensive research is now being conducted in the liquid phase [50,51]. In addition to a lower catalyst deactivation rate, liquid phase reaction also reduces the number of side-products. The drawbacks are, of course, lower activity but also abrasion problems with the catalyst. The catalyst is preferably stabilized as a powder in a silica matrix (Ludox R) [51]. The catalyst most often encountered in the patent literature is a Cu-Cr with a promoter such as Ba or Mn. The catalyst is also preferably doped with Na or K and pretreated very carefully in a reducing atmosphere [52]. 

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