Nylon 4,I: an amorphous polyamide

The melt polymerization of nylon 4, I was studied, starting with nylon-salt and nylon prepolymers (ηinh=0.25). With nylon-salt only low molecular polymers were obtained, while with prepolymers the inherent viscosity could be raised to 0.77 (3h, 270°, vac.). The cyclization of tetra methylene diamine to pyrrolidine seem to be the major disturbing factor. The polymer is glassy and could easily be melt pressed. The torsion modulus of the material at 20°C was high (1.8 109 Pa) and remained high to near its Tg (138°C)

A small scale regularly packed circulating fluidized bed. Part II: Mass transfer

The underlying objective of the present study is to increase gas¿solids contact in a circulating fluidized bed by the introduction of obstacles in the riser portion. The presence of such obstacles leads to suppression of radial inhomogeneities in the solids mass flux and concentration, and break-up of solids clusters. At ambient conditions, gas¿solids mass transfer was investigated for cocurrent upward flow of air and microsize solid particles (FCC, 70 ¿m diameter) over a regularly structured inert packing introduced into the riser part of a circulating fluidized bed unit. The packed section has a height of 0.48 m, a cross-sectional area of 0.06 × O.06 m2, and contains regularly stacked 0.01 m diameter Perspex bars as the obstacles meant to enhance the gas¿solids contact. Gas mass fluxes used were 1.4 and 2.7 kg m¿2 s¿1. Solids mass fluxes were varied in the range 0Gs 12 kg m¿2 s¿1. Experimental mass transfer data were obtained by applying the method of adsorption of naphthalene vapor on FCC particles. A conservative estimate of the apparent gas¿solids mass transfer coefficient kg* could be derived from the naphthalene vapor concentration profile along the packed section on the basis of a plug-flow-model interpretation, while assuming single-particle behaviour and neglecting intraparticle diffusion effects. Such kg* values appear to increase with increasing gas mass flux, but decrease with increasing solids mass flux (and consequently increasing solids volume fraction) probably due to the corresponding increase in particle shielding. Comparison of the present results with available literature data for similar solid materials suggests that the effect of the packing inserted into the CFB is significant: the Sherwood numbers derived from the present study are relatively high

A small scale regularly packed circulating fluidized bed. Part I: Hydrodynamics.

The present investigation is based on the idea of intensifying the gas¿solids contact in a circulating fluidized bed by introducing obstacles into it. Such obstacles may effectively suppress radial inhomogeneities in the solids flux and concentration, increase the dynamic solids hold-up, and break up solids clusters. This article (Part I) deals with the hydrodynamics (pressure drop and solids hold-up) investigated at ambient conditions, for cocurrent upward flow of air and microsize solid particles (FCC, 70 µm diameter) over a regularly structured inert packing introduced into the riser part of a circulating fluidized bed unit. The packed section has a height of 0.48 m, a cross-sectional area of 0.06 × 0.06 m2 and contains regularly-stacked 0.01 m diameter Perspex bars as the obstacles meant to enhance the gas¿solids contact. Slide-valves mounted above and below the packed section can be used to trap the solids inventory and determine the (dynamic) solids hold-up. Gas and solids mass fluxes have been varied in the range of 0.7 < Gg < 4.4 and O < Gs < 15 kg m-2s-2, respectively. Part II will report on the results of gas¿solids mass transfer measurements, which have been carried out in the same set-up at comparable experimental conditions. Results of this work show that: (i) the pressure gradient over the packed section increases linearly with increasing solids mass flux, but faster than linearly with increasing applied gas mass flux, (ii) the dynamic solids volume fraction can be described quite well by the correlation ß dyn = 0.0084 GsGg-1.22 for almost the entire range of applied gas and solids mass fluxes, (iii) the value for the solids friction factor derived for the gas flux range 0.7 < Gg < 3.7 kg (m-2s-1) varies from 1.4 to 2.5 and is linear with the solids volume fraction. These fs values are about 2 to 3 decades higher than those obtained from fs correlations derived for dilute-phase pneumatic conveying lines operated under the same experimental conditions

Economic comparison of reactive distillation (RD) to a benchmark conventional flowsheet:Regions of RD applicability and trends in column design

A novel methodology for the techno-economic assessment of Reactive Distillation (RD) is presented. The developed methodology benchmarks reactive distillation (RD) to a conventional reactor + distillation train flowsheet (R+D) on a cost-optimized basis, with the optimization being performed on the process unit level (reactor sizing, number of stages, feed point(s)) and the internals level (reactive tray design). This methodology is applied to the ideal quaternary system A+B↔C+D with the conventional boiling point order of TC &lt; TA &lt; TB &lt; TD (αAD = 4, αBD = 2, αCD = 8). From this pool of data, a regime map of RD vs. R+D is established in which the attractive regions of either flowsheet option are identified in terms of the chemical reaction rate and chemical equilibrium. It is found that RD can arise as the cost optimal option for a large range of residence time requirements by virtue of overcoming the external recycle requirements of R+D. This is achieved through optimized reactive tray design. Contrary to conventional distillation design practices, it was found that the preferred use of bubble-cap trays over sieve trays to allow elevated weir heights and designing the column diameter below 80% of flooding become relevant design choices when accommodating high liquid holdup.</p

The industrial production of dimethyl carbonate from methanol and carbon dioxide

This work discusses the design of a dimethyl carbonate (DMC) production plant based on methanol and CO2 as feed materials, which are a cheap and environment-friendly feedstock. DMC is a good alternative for methyl-tert-butyl ether (MTBE) as a fuel oxygenating agent, due to its low toxicity and fast biodegradability. Based on the MTBE demand of a general gasoline plant, the annual production capacity of the process design is stipulated to be 86 kt DMC, with a purity of 99 wt%. Three routes are proposed to form DMC: 1) direct synthesis from methanol and CO2, 2) reaction of CO2 and ammonia to urea, which can be converted to DMC with methanol, 3) reaction of ethylene oxide with CO2 to a cyclic carbonate, which can be converted to DMC by transesterification with methanol. From a black box cost analysis based on raw material prices, it is concluded that the ethylene oxide route is the least profitable. Because of higher single-pass conversions found in literature, smaller recycles and easier separations, it is concluded that the urea route would be the most feasible. The required process functions for the urea route have been determined in the conceptual design phase. A detailed design of the most important process operations is made and an overall technical and economic evaluation of the process has been carried out. In the first step of this DMC synthesis, urea is produced from carbon dioxide and ammonia with the ACES21 process. After separation and purification steps, urea is fed to a reactor with methanol (150 °C, 20 bar), where methyl carbamate (MC), an intermediate of DMC production, and ammonia are formed in the absence of a catalyst. Subsequently, MC and methanol are converted to DMC and ammonia (190 °C, 40 bar) over a ZnO-Al2O3 catalyst in a fixed-bed reactor. Methanol and DMC form an azeotrope; extractive distillation with methyl isobutyl ketone (MIBK) as entrainer is used to separate the azeotropic mixture. The reactor model for the reaction towards DMC based on kinetic rate expressions, showed that a long residence time (>10 h) and a relatively high MeOH:MC molar feed ratio of 6 are required to achieve reasonable single-pass conversions (15 %). This resulted however in an unrealistically large reactor volume and a large methanol load on the process. A feasibility study was done in order to improve the performance of the process. It was calculated that with a MeOH:MC ratio of 2 and a single-pass conversion of MC of 30 % the process would become technically feasible; the reactor volume decreased from 5,000 m3 to 600 m3 and the energy consumption of the process was decreased from 238 MW to 50 MW. A Pinch analysis showed that maximally 6 MW could be saved with heat integration, which corresponds to approximately 2 M$/y savings on energy costs. To produce 86 kt/y of DMC, the required amounts of raw materials are 80 kt/y of methanol and 58 kt/y of CO2, which results in an overall DMC yield from methanol of 38$. Economic feasibility depends on the DMC selling price. A price range between 800 and 1,100 $/t was assumed. For 800$/t it is not possible to repay the capital investment within an assumed lifetime of 10 years and the process would therefore not be profitable. The break-even point is at 845 $/t. For a selling price of 1,100$/t the gross profit becomes 22 M$/y, with a payback period of 3 years and a return on investment of 20 %

Liquid organic hydrogen carriers:Process design and economic analysis for manufacturing N-ethylcarbazole

This paper revisits the economics of manufacturing N-ethylcarbazole (NEC), a strong candidate for large-scale liquid organic hydrogen carrier (LOHC) supply chains, because of its high H2 storage capacity (6 wt%), selective hydrogenation and dehydrogenation reactions, and favorable reaction enthalpy and reaction temperatures compared to other LOHC systems. Two different process routes for producing NEC from industrial chemicals are selected out of 10 possible options: one using aniline and the other using cyclohexanone and nitrobenzene as feedstock. The required capital and operational costs are estimated to determine a NEC break-even cost for a capacity of 225 ktpa. NEC break-even costs of $3.0 and$2.6 per kg LOHC are found for the routes. This is significantly less than the $40/kg cost that has generally been reported in literature for NEC, thus improving the economic viability of using NEC as LOHC. The total fixed capital costs are estimated to be$200 MM and $250 MM. Furthermore, the prices of the feedstock show the largest influence (76$0.77–$0.90 per kg H2 for a 60-day roundtrip and$0.09–$0.10 per kg H2 for a 7-day roundtrip. It is important to note that both routes rely heavily on laboratory scale data and the corresponding assumptions that stem from this limitation. Therefore, this research can serve as a guide to future experimental studies into validating the key assumptions made for this analysis.</p

Pervaporative separation and intensification of downstream recovery of acetone-butanol-ethanol (ABE)

The feasibility of pervaporative concentration of organic compounds from an ABE mixture to reduce the energy consumption of a downstream recovery unit was investigated. Firstly, an experimental investigation was done, using a polydimethylsiloxane (PDMS) membrane and a model solution of ABE as the feed. Different operating temperatures where investigated, with 40 °C showing the most favourable results. Secondly, the experimental results were utilised as the input for process simulations using Aspen Plus. Two ABE separation schemes were studied, one consisting of only distillation (conventional process) and one with an upstream pervaporation unit followed by an alternative distillation scheme. For the proposed pervaporative scheme, the butanol concentration after pervaporation was high enough so that it could be concentrated further at the beginning of the separation train through a liquid-liquid separation. The results of the simulations indicated that the conventional scheme was the most energy intensive and that the integration with an upstream pervaporation unit decreased energy consumption with 53%. The energy requirement for the distillation scheme was 33.3 MJ kg−1 butanol, while that of the pervaporation-distillation scheme was 15.7 MJ kg-1 butanol

Evolution patterns and family relations in G–S reactors

Reactor selection strategies for gas–solid (G–S) heterogeneously catalysed processes can be based on the requirements of the desired process and the properties of the reactions and catalysts involved. Ultimately a reactor selection will nearly always be grounded on existing or emerging reactor types slightly modified for adaptation to the specific chemical process. This procedure results in radiation of different reactor modifications from the archetypes towards niche applications. It is shown that this process has a lot of resemblance with the evolution process of animal species. The G–S heterogeneous catalytic reactors can be classified into three or four families. They are presented as adaptations from only three archetypes: packed bed, fluid bed and barrier wall.\ud \ud The properties of these reactors and their family members are discussed. Examples are given of a few relatively new variants and of competition of very different reactors for the same application niche. The classification system can be used as a means for the creation of new reactors and if extended with a database or knowledge system it can facilitate reactor selection. Similar classifications can be set-up for other types of chemical reactors like G–L and G–L–S reactors