Catalyst beds industrial data

Table 19.3 (after this chapter s problems) gives industrial after-H2S04-making catalyst bed operating data. 

In this section we refer to the same industrial reactor as in 7.5.1, with its data given on p. 508. Further reactor specifications and catalyst-bed properties of this plant are as follows. 

Reactor Model. The design of an industrial packed-bed reactor requires a reactor model as well as the chemical and the heat and mass transfer parameters of the catalyst bed - gas stream system. Since these parameters are model-specific, it seemed advisable to employ a continuum model for the reactor calculation. This is the only model to date for which the literature contains consistent data for calculating heat and mass transfer parameters (5,6,7). This model in its... 

Based on the results of Dalla Betta and co-workers, it is clear that the steady-state activity of a completely sulfur-poisoned Ni or Ru methanation catalyst is 102-104 times lower than that of the fresh catalyst. However, a typical industrial methanation process would more probably involve a catalyst only partly poisoned by sulfur. Bartholomew and co-workers (23, 113, 157) attempted to assess how sulfur poisoning of only a portion of the catalyst would affect its activity/selectivity properties in fixed-bed and fluidized-bed reactors. Data in Table XII show the effects on specific activity and product distribution of partially presulfided Co/A1203 and Ni/Al203 catalysts in a fixed bed. Catalysts were presulfided with 10 ppm H2S at 725 K, and reaction was carried out with sulfur-free feedgas. Corresponding data are listed in Table XIII for catalysts partially presulfided and then studied in a fluidized-bed reactor under the same conditions. The decrease in H2 uptake... 

Third, there may be a concentration gradient of reactants and products along the length of the catalyst bed. If the structure of the catalyst depends upon the composition of the gas phase, then an average of the various structures will be measured. There is little discussion of this topic in the literature of XAFS spectroscopy of working catalysts. An extreme example of structural variations within a sample is discussed in Section 6, where there is a discussion of XAFS spatially resolved spectra recorded to allow direct observation of the axial distribution of phases present. If the XAFS data are not measured with spatial resolution, then it is recommended that XAFS data be measured under differential conversion conditions. However, if the aim of the experiment is to relate the catalyst structure directly to that in some industrial catalytic processes, then differential conversion conditions will only reflect the structure of the catalyst at the inlet of the bed. To learn about the structure of the catalyst near the outlet of the bed, the reaction has to be conducted at high conversions. If it is anticipated that this operation will lead to variations in the catalyst structure along the bed, then the feed to the micro-reactor should be one that mimics the concentration of reactants toward the downstream end of the bed (i.e., products should be added to the reactants). 

Industrial data (Table 7.2) suggest that daisy rings and rings are equally favored. Pellets are only used to ensure well distributed gas flow through catalyst beds in low gas velocity converters (Topsoe, 2004). 

Data obtained on trickle-flow processes obtained in microflow tests with diluted catalyst beds are not only reproducible, but are also meaningful for industrial practice. This is demonstrated in Table X by the direct comparison of microflow test results with data from an industrial gasoil hydrodesulfurization unit. 

In the mechanistic studies conducted by Horn et al.,64 methane CPO reactions were run under as close to industrial conditions as possible. A high-resolution (=300/rm) spatial profiling technique was developed to measure the temperature and species profiles along the centerline of the catalyst bed. The unprecedented high-resolution data make it possible to accurately speculate reactions occurring along the catalyst bed (10mm) in a very short contact time. It was revealed that for Rh-supported CPO catalysts, H2 and CO are formed in both the initial 0-2-mm... 

Table 11.14 gives tcchnico-economic data concerning the two principal processes for manufacturing acrylonitrile currently industrialized, and which involve the ammoxi-dation of propylene in a fluidized bed or a fixed catalyst bed. 

See Table 19.3 (end of this chapter) for industrial after H2SO4 making catalyst bed data. 

The steady- and non steady-state simulation of an industrial ammonia converter is presented. The reactor includes two adiabatic radial-flow catalyst beds in series. An interbed (gas-gas) heat exchanger is used to preheat the feed stream. The steady-state results showed good agreement with plant data. The influence of different disturbances (feed composition and temperature, reactor pressure) on the dynamic evolution of the main variables is analysed. The open-loop and closed-loop operation is compared from the standpoint of the reactor stability. 

In trickle-bed microreactors (TBMR), it is frequently used a commercial catalyst sample and real feedstocks, however, the length of the catalyst bed and hence the reactor length to catalyst particle diameter ratio are low as compared to commercial reactors. In addition, low liquid velocities are used in order to match the liquid hourly space velocities (LHSV) of industrial units. These differences cause number of problems in testing catalyst having commercially applied size and shape, such as poor wetting of catalyst, wall effect, axial dispersion, maldistribution, and the data obtained in such TBMR may not be reliable [1,2],... 

Simulation of tubular steam reformers and a comparison with industrial data are shown in many references, such as [250], In most cases the simulations are based on measured outer tube-wall temperatures. In [181] a basic furnace model is used, whereas in [525] a radiation model similar to the one in Section 3.3.6 is used. In both cases catalyst effectiveness factor profiles are shown. Similar simulations using the combined two-dimensional fixed-bed reactor, and the furnace and catalyst particle models described in the previous chapters are shown below using the operating conditions and geometry for the simple steam reforming furnace in the hydrogen plant. Examples 1.3, 2.1 and 3.2. Similar to [181] and [525], the intrinsic kinetic expressions used are the Xu and Froment expressions [525] from Section 3.5.2, but with the parameters from [541]. 

Figure 8.7 shows industrial catalyst bed input gas temperamres (Tables 7.2-7.6). There is considerable spread in the data because various catalysts are being used. 

Figure 9.6 Double contact H2SO4 making flowsheet. The two absorption towers are notable. The left half of the flowsheet oxidizes most of the S02-in-feed-gas and makes the product SO3 into strengthened sulfiiric acid. It makes about 95% of the plant s new H2SO4. The right half of the flowsheet oxidizes almost all the remaining SO2 and makes its product SO3 into strengthened sulfuric acid. The final exit gas is very dilute in SO2 and SO3. Industrially, all the catalyst beds are in the same converter (Fig. 7.7). Table 23.2 gives industrial final H2SO4 making data.

The SO2 -l- O.5O2 —> SO3 reaction kinetics explain why equilibrium is not achieved in each catalyst bed. The kinetics are complex. Many studies have been carried out (Davenport and King, 2006), but the best information resides with catalyst manufacturers such as BASF, Haldor Topspe, and MECS. A literature review, along with industrial data, indicates the following ... 

Some modes of heat transfer to stirred tank reacdors are shown in Fig. 23-1 and to packed bed reactors in Fig. 23-2. Temperature and composition profiles of some processes are shown in Fig. 23-3. Operating data, catalysts, and reaction times are stated for a number of industrial reaction processes in Table 23-1. 

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