Capital cost Catalyst

Future Methanol Processes. The process route for methanol synthesis has remained basically unchanged since its inception by BASF in 1923. The principal developments have been in catalyst formulation to increase productivity and selectivity, and in process plant integration to improve output and energy efficiency while decreasing capital cost. 

The same four operating steps are used with the complex batch reactor as with the simple batch reactor. The powerhil capabiUties of the complex batch reactor offset their relatively high capital cost. These reactors can operate at phenol to alkene mole ratios from 0.3 to 1 and up. This abiUty is achieved by designing for positive pressure operation, typically 200 to 2000 kPa (30 to 300 psig), and for the use of highly selective catalysts. Because these reactors can operate at low phenol to alkene mole ratios, they are ideal for production of di- and trialkylphenols. 

Base Metal Catalyst - An alternate to a noble metal catalyst is a base metal catalyst. A base metal catalyst can be deposited on a monolithic substrate or is available as a pellet. These pellets are normally extruded and hence are 100% catalyst rather than deposition on a substrate. A benefit of base metal extruded catalyst is that if any poisons are present in the process stream, a deposition of the poisons on the surface of the catalyst occurs. Depending on the type of contaminant, it can frequently be washed away with water. When it is washed, abraded, or atritted, the outer surface is removed and subsequently a new catalyst surface is exposed. Hence, the catalyst can be regenerated. Noble metal catalyst can also be regenerated but the process is more expensive. A noble metal catalyst, depending on the operation, will typically last 30,000 hours. As a rule of thumb, a single shift operation of 40 hours a week, 50 weeks a year results in a total of 2,000 hours per year. Hence, the catalyst might have a 15 year life expectancy. From a cost factor, a typical rule of thumb is that a catalyst might be 10%-15% of the overall capital cost of the equipment. 

Determination of the actual cost of a hydrogenation process is difficult. Among the factors entering into the determination are catalyst cost, catalyst life, cost of materials, capital investment, actual yield, space-time yield, and purification costs, Considerable data are needed to make an accurate evaluation. 

Even in a simple hydrogen fuel cell system, capital cost reduction requires improvements in many diverse areas, such as catalyst loadings, air pressuriza-... 

The rate of a chemical reaction is influenced by pressure, temperature, concentration of reactants, kinetic factors such as agitation, and the presence of a catalyst. Since the viability of a plant depends not only on reaction efficiencies but also on the capital cost factor and the cost of maintenance, it may be more economic to alter a process variable in order that a less expensive material of construction can be used. The flexibility which the process designer has in this respect depends on how sensitive the reaction efficiency is to a change in the variable of concern to the materials engineer. 

Of these, fixed-bed adiabatic reactors are the cheapest in terms of capital cost. Tubular reactors are more expensive than fixed-bed adiabatic reactors, with the highest capital costs associated with moving and fluidized beds. The choice of reactor configuration for reactions involving a solid supported catalyst is often dominated by the deactivation characteristics of the catalyst. 

The advantage of catalytic thermal oxidation is that the lower temperature of operation can lead to fuel savings (although effective heat recovery without a catalyst can offset this advantage). The major disadvantages of catalytic thermal oxidation are that the catalyst needs to be replaced every two to four years and the capital cost tends to be higher than thermal oxidation without a catalyst. Catalytic thermal oxidation also tends to increase the pressure drop through the system. 

Catalyst cost constitutes 15-20% of the capital cost of an SCR unit therefore, it is essential to operate at temperatures as high as possible to maximize space velocity and thus minimize catalyst volume. At the same time, it is necessary to minimize the rate of oxidation of S02 to S03, which is more temperature sensitive than the SCR reaction. The optimum operating temperature for the SCR process using titanium and vanadium oxide catalysts is about 38CM180oC. Most installations use an economizer bypass to provide flue gas to the reactors at the desired temperature during periods when flue gas temperatures are low, such as low-load operation. 

The choice of cyclone modification, from an operating point of view, becomes a balance of incremental profit from increased conversion, versus catalyst makeup charges, and from a capital cost point of view, the price of either of the cyclone modifications, which must be depreciated. In many instances, there is an additional background time element, involving ongoing development of more attrition resistant and/or active catalyst. 

Another aspect of gas recycle that would need to be considered is the degree of catalyst solution expansion that would result from very high gas recycle flows. The catalyst solution might be blown out of the reactor, or if a very tall reactor were built, it would suffer from excessive capital costs for the large containment needed for the gas-expanded catalyst. 

At these lower conversions, the previously described multi-bed down-flow reactor may not be the best design. At the higher space velocities it is more expedient to have the liquid in up-flow, and the reduced penalty from axial (particularly) forward mixing means also that the necessary conversion can frequently be achieved in a single catalyst bed. Thus, the reactor design is considerably simplified, with an associated reduction in the capital cost for a given catalyst volume. The development of the design for an up-flow-packed bed reactor for the decomposition of sodium hypochlorite is described elsewhere [6] and will not be considered in detail here. 

For both low temperature electrolysers, the biggest gain in efficiency is to be expected from an improvement in Balance of Plant components, taking into account the big gap between cell efficiency (80-90%) and system efficiency (50-60%). In the case of SPE electrolysers, catalytic research should therefore be directed to making the catalysts more tolerant to contaminants. For alkaline electrolysers, in addition to this, more active electrodes could lower capital costs. 

In 1994, Rayox developers generated cost estimates for the technology based on bench-scale studies. Using a proposed cleanup site in Canada as a model, researchers compared these estimates with the costs of using an air stripper/liquid carbon/catalytic oxidizer (air/carbon) option. Results indicated that UV/peroxide treatment, with or without an iron catalyst, was found to have an estimated capital cost equal to the air/carbon option at the site (D12302U). 

It is implicit in reaction 9.4 that the equilibrium yield of ammonia is favored by high pressures and low temperatures (Table 9.1). However, compromises must be made, as the capital cost of high pressure equipment is high and the rate of reaction at low temperatures is slow, even when a catalyst is used. In practice, Haber plants are usually operated at 80 to 350 bars and at 400 to 540 °C, and several passes are made through the converter. The catalyst (Section 6.2) is typically finely divided iron (supplied as magnetite, Fe304 which is reduced by the H2) with a KOH promoter on a support of refractory metallic oxide. The upper temperature limit is set by the tendency of the catalyst to sinter above 540 °C. To increase the yield, the gases may be cooled as they approach equilibrium. 

Another factor which contributes most to the high cost of the alkylation process is the necessity of having a large excess of isobutane in the reaction zone. This requirement results in increased capital costs for over-sized reactors, settlers, fractionators, and accessories, as well as in increased operating costs of this equipment. Process developments which would allow satisfactory operations at considerably lower isobutane-to-olefin ratios would help reduce these costs. Such developments might again involve improvements in reactors or in catalysts. 

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