Preheating phase

A fuel cell is an electrochemical device for converting the chentical energy of a fuel directly into electricity, without the efficiency losses of conventional combustion processes which are constrained by the Carnot limit. In principle, the conversion efficiency is 100% for FCs, while it is around 35-40% for the Carnot corrversion at the temperatures itsed. In real cases, efficiencies of FCs are lowered by the requirements of other plant componerrts, such as reforming arrd preheating phases as well as irreversibilities affecting the process. 

This preheating phase is critical for the industry since it can have a major impact on the cell lifespan (Tessier et al., 2011 Tessier et al., 2010) and was thus closely investigated over the past years through numerical simulations (D Amours et al., 2003 Marceau et al., 2011 Richard et al., 2006 Sun et al., 2004). It is therefore required to develop constitutive laws for the three aforementioned materials to feed numerical models. The targeted materials are briefly presented below. 

Furthermore, 60—100 L (14—24 gal) oil, having sulfur content below 0.4 wt %, could be recovered per metric ton coal from pyrolysis at 427—517°C. The recovered oil was suitable as low sulfur fuel. Figure 15 is a flow sheet of the Rocky Flats pilot plant. Coal is fed from hoppers to a dilute-phase, fluid-bed preheater and transported to a pyrolysis dmm, where it is contacted by hot ceramic balls. Pyrolysis dmm effluent is passed over a trommel screen that permits char product to fall through. Product char is thereafter cooled and sent to storage. The ceramic balls are recycled and pyrolysis vapors are condensed and fractionated. 

Naphtha desulfurization is conducted in the vapor phase as described for natural gas. Raw naphtha is preheated and vaporized in a separate furnace. If the sulfur content of the naphtha is very high, after Co—Mo hydrotreating, the naphtha is condensed, H2S is stripped out, and the residual H2S is adsorbed on ZnO. The primary reformer operates at conditions similar to those used with natural gas feed. The nickel catalyst, however, requires a promoter such as potassium in order to avoid carbon deposition at the practical levels of steam-to-carbon ratios of 3.5—5.0. Deposition of carbon from hydrocarbons cracking on the particles of the catalyst reduces the activity of the catalyst for the reforming and results in local uneven heating of the reformer tubes because the firing heat is not removed by the reforming reaction. 

In the heating and cracking phase, preheated hydrocarbons leaving the atomizer are intimately contacted with the steam-preheated oxygen mixture. The atomized hydrocarbon is heated and vaporized by back radiation from the flame front and the reactor walls. Some cracking to carbon, methane, and hydrocarbon radicals occurs during this brief phase. 

A Hquid-phase variation of the direct hydration was developed by Tokuyama Soda (78). The disadvantages of the gas-phase processes are largely avoided by employing a weakly acidic aqueous catalyst solution of a siHcotungstate (82). Preheated propylene, water, and recycled aqueous catalyst solution are pressurized and fed into a reaction chamber where they react in the Hquid state at 270°C and 20.3 MPa (200 atm) and form aqueous isopropyl alcohol. Propylene conversions of 60—70% per pass are obtained, and selectivity to isopropyl alcohol is 98—99 mol % of converted propylene. The catalyst is recycled and requites Htde replenishment compared to other processes. Corrosion and environmental problems are also minimized because the catalyst is a weak acid and because the system is completely closed. On account of the low gas recycle ratio, regular commercial propylene of 95% purity can be used as feedstock. 

The OLEFLEX process uses multiple side-by-side, radial flow, moving-bed reactors connected in series. The heat of reaction is suppHed by preheated feed and interstage heaters. The gas-phase reaction is carried out over a catalyst, platinum supported over alumina, under very near isothermal conditions. The first commercial installation of this technology, having an annual capacity of 100,000 t, was made in 1990 by the National Petrochemical Corporation in Thailand. A second unit, at 245,000 t capacity, has been built in South Korea by the ISU Chemical Company (70). 

Gas Phase. The gas-phase methanol hydrochlorination process is used more in Europe and Japan than in the United States, though there is a considerable body of Hterature available. The process is typicaHy carried out as foHows vaporized methanol and hydrogen chloride, mixed in equimolar proportions, are preheated to 180—200°C. Reaction occurs on passage through a converter packed with 1.68—2.38 mm (8—12 mesh) alumina gel at ca 350°C. The product gas is cooled, water-scmbbed, and Hquefied. Conversions of over 95% of the methanol are commonly obtained. Garnma-alurnina has been used as a catalyst at 295—340°C to obtain 97.8% yields of methyl chloride (25). Other catalysts may be used, eg, cuprous or zinc chloride on active alumina, carbon, sHica, or pumice (26—30) sHica—aluminas (31,32) zeoHtes (33) attapulgus clay (34) or carbon (35,36). Space velocities of up to 300 h , with volumes of gas at STP per hour per volume catalyst space, are employed. 

Elastomer-plastic blends without vulcanization were prepared either in a two roll mill or Banbury mixer. Depending on the nature of plastic and rubber the mixing temperature was changed. Usually the plastic was fed into the two roll mill or an internal mixer after preheating the mixer to a temperature above the melting temperature of the plastic phase. The plastic phase was then added and the required melt viscosity was attained by applying a mechanical shear. The rubber phase was then added and the mixture was then melt mixed for an additional 1 to 3 min when other rubber additives, such as filler, activator, and lubricants or softeners, were added. Mixing was then carried out with controlled shear rate... 

The hexane solvent was removed from a solution of DIB AL (22 mmol, 1 Min hexane) at reduced pressure and at ambient temperature, and ether (10 ml) was introduced. l-Cyclohexyl-2-trimethylsilylethyne (20 mmol) was added at such a rate as to maintain ambient temperature within the reaction, and. after 15 min, the reaction flask was placed in a preheated (40 °C) bath for I h. The resulting clear solution was transferred by means of a double-ended syringe to a vigorously stirred cold solution of HC1 (50 ml, 10%). The flask was rinsed with ether (20 ml), and the mixture was stirred until the resulting phases were almost clear. The layers were separated, and the aqueous layer was extracted with ether (40 ml). The combined organic extracts were washed successively with dilute HC1 (20%), saturated sodium hydrogen carbonate solution and brine, and dried. 

After leaving the reactor the reaction mixture is passed to a settling tank where the denser HF is deposited in the lower phase. The organic phase is mixed gently with HF the HF phase contains tar components and traces of benzene. From the HF phase a side stream is refined. This side stream is heated in a preheater, partially vaporized, and separated into two components in a distillation column HF and benzene are distilled over the top while tar components are taken away at bottom. The top product is condensed, cooled, and collected in a settle tank. The bottom product is neutralized using potassium... 

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