Development of a Process for Higher Alcohol Production via Synthesis

3896

Ind. Eng. Chem. Res. 1998, 37, 3896-3908

Development of a Process for Higher Alcohol Production via Synthesis Gas Alessandra Beretta,† Emilio Micheli,‡ Lorenzo Tagliabue,‡ and Enrico Tronconi*,† Dipartimento di Chimica Industriale e Ingegneria Chimica, Politecnico di Milano, Piazza Leonardo da Vinci 32, 20100 Milano, Italy, and Snamprogetti SpA, Research Laboratories, via Maritano 26, 20097 San Donato Milanese, Italy

This paper reports work at the laboratory and pilot scale concerning the design of a process for the production of valuable higher oxygenates from synthesis gas. The performance of a Snamprogetti high-temperature catalyst for the synthesis of methanol and higher alcohols was extensively studied, and a lumped kinetic model was developed; this accounts for the product distribution over the range of interest of temperature, pressure, and space velocity and includes the effect of synthesis gas composition. The model was applied to the preliminary simulation of a multistage adiabatic reactor. A single-stage adiabatic pilot reactor was then realized and operated. It was verified that higher alcohol synthesis can be carried out without the risk of uncontrollable heat buildup by properly tuning the amount of methanol in the feed stream. Introduction The research on the synthesis of methanol and higher oxygenates from synthesis gas dates back to the beginning of the century. Yet, higher alcohol synthesis (HAS) still remains in the field of research and development efforts as few demonstrative plants have been realized (Paggini et al., 1986; Ohno et al., 1986) but no commercial applications are known. As compared to methanol synthesis, the rationalization of the catalyst constituents and the operating conditions which favor the formation of higher alcohols as well as the investigation of the reaction mechanism and kinetics have received much less attention, and mostly in recent years (Hindermann et al., 1993; Forzatti et al., 1991). The hydrogenation of carbon monoxide to oxygenates with more than two carbon atoms is a very complex reaction; it involves a huge variety of intermediates and products which differ in chemical nature (primary and secondary alcohols, aldehydes and ketones, carboxylic acids, esthers, ethers, both saturated and unsaturated hydrocarbons, waxes, multifunctional species), carbon atom number, and molecular structure (linear and single- and multibranched). The reaction pathways that link these species are then specific of the catalytic systems. CO-insertion mechanisms dominate over the modified Fischer-Tropsch catalysts (producing mainly linear primary alcohols with Anderson-Schultz-Florytype distributions) (Hindermann et al., 1993), aldol-type condensation mechanisms of Cn with C1 species prevail over the low-temperature Cu-based catalysts (where high productions of ethanol, propanol, and isobutyl alcohol accompany the synthesis of methanol) (Nunan et al., 1988; Smith et al., 1991), and Cn + C1 and Cn + Cm condensations and ketonizations are active over high-temperature alkali-modified methanol catalysts (which give high selectivities and yields to isobutyl alcohol and other 2-methyl primary alcohols) (Forzatti et al., 1991; Tronconi et al., 1992). * To whom correspondence should be addressed. Fax: +39 02 70638173. E-mail: [email protected]. † Politecnico di Milano. ‡ Snamprogetti SpA.

The performances of the number of catalysts reported in the literature, however, are all below the targets of feasibility for commercial processes, if the interest relies either on the whole pool of higher oxygenates for gasoline blending (a proposal in the 1980s) or on single alcohols regarded as intermediates for further transformations, as more recently proposed. At present, isobutyl alcohol is believed to be a potential precursor for methyl tert-butyl ether production (Underwood and Schaub, 1993; Zhou, 1994; Sanfilippo et al., 1998). Beretta et al. (1996a) have shown that the production of isobutyl alcohol can be significantly increased by coupling the catalytic properties of a Cs-doped, Cucontaining catalyst and a Cs-doped, Cu-free catalyst in a double-bed configuration. In this work, the design of a process for the production of higher alcohols over a high-temperature-modified methanol synthesis catalyst is addressed. Studies in Snamprogetti laboratories (Antonelli and Cornaro, 1995) led to the development of a zinc-chromite catalyst which produced high yields of isobutyl alcohol and other branched alcohols. In the following, a kinetic study over the catalyst powders and the development of a simplified model are first presented. The construction, the operation, and the model analysis of a pilot-scale adiabatic higher alcohol synthesis unit are then reported. The novelty of this work is represented by the pursuit of experimental and theoretical standpoints about the industrial feasibility of an adiabatic reactor for higher alcohol synthesis. No similar attempts have been published in the literature, yet. While the bulk of the literature data refer to conventional bench-scale studies, rather alarming indications were given by Stiles et al. (1991) about the industrial realization of the process; possible uncontrolled exothermicity of the reaction related to the undesired production of methane and higher hydrocarbons were, in fact, suggested. The results herein presented are thus especially promising in relation to both the accomplishment of high productivities to isobutyl alcohol and the realization of an industrial plant.

S0888-5885(97)00733-1 CCC: $15.00 © 1998 American Chemical Society Published on Web 09/16/1998

Ind. Eng. Chem. Res., Vol. 37, No. 10, 1998 3897

Experimental Section Catalyst Preparation. The catalyst used in the present study was a K-promoted Zn/Cr/O system with a Zn/Cr mole ratio equal to 0.75 (Antonelli and Cornaro, 1995). A 0.6 L solution of zinc and chromium nitrates (1.5 total metal mol/L of distilled water) and 0.9 L of a 2.0 M (NH4)2CO3 solution were added by high-precision pumps to an initial volume of 0.3 L of distilled water. Coprecipitation was performed at ambient temperature under vigorous stirring, within 1 h. The precipitate was aged by keeping the suspension stirring for an additional 30 min. The suspension was then washed and filtered repeatedly until a complete removal of nitrates was achieved. The precipitate was finally redispersed in 1 L of distilled water and spray-dried. Drying was carried out at 383 K for 12 h. Calcination was then performed at 673 K for 4 h under flowing N2. The calcined catalyst was impregnated with an aqueous solution of potassium carbonate, for a final K loading of 3 wt %. After impregnation the catalyst was dried again. The catalyst was used in the form of 14-20 mesh particles for the kinetic study and in the form of 3 × 3 mm cylindrical pellets in the adiabatic reactor. The catalyst showed a high stability; it was tested for a period of 2500 h under the most severe conditions (T ) 693 K, P ) 18.0 MPa) without occurrence of appreciable deactivation. The BET surface area amounted to 165 m2/g for the fresh catalyst and to 140 m2/g for the spent catalyst. Kinetic Study: Testing Unit. Pure CO, H2, and CO2 were fed from cylinders, premixed, and compressed upstream from the reactor. The CO stream was purified by passing through a guard bed of inert material kept at 723-773 K inside a Cu tube; iron carbonyls were thus decomposed. The testing apparatus used in the present study is schematically represented in Figure 1. A mass flowmeter and a pneumatic valve measured and regulated the feed stream. A pneumatic valve located before the outlet vent line controlled the total pressure. The reactor consisted of a stainless steel Cu-lined tube with an internal diameter of 18 mm. It was heated by a fluidized sand bath. A thermocouple well was located inside the reactor, and a sliding thermocouple was used to measure the catalyst axial temperature profile. A total of 4-5 g of catalyst particles was usually diluted with inert material and loaded inside the reactor. The catalyst was activated using the same synthesis gas mixture (gas hourly space velocity (GHSV) ) 20 000 L(NTP)/kg of catalyst/h) at a total pressure of 2 MPa by heating the reactor up to 693 K for 2 h. The reactor temperature was then lowered to 613 K; the rig was subsequently pressurized and heated again to the operating temperature. Downstream from the reactor, the liquid product was condensed at the system pressure in a separator, watercooled at 288 K. The flow rate of the residual gas stream was measured by a Symbrunt AB-1 wet flowmeter; a portion of the gas stream (1-2 L(NTP)/h), measured by a rotameter, was sent to a HP 5880 gas chromatograph (GC) for on-line analysis. A Porapak Q column using H2 as the carrier and connected to a thermoconductivity detector (TCD) was used for the analysis of CO, CO2, and CH4. A Porapak Q + Porapak T column connected to a flame ionization detector was used to analyze the C1-C4 hydrocarbons, dimethyl

ether, and the methanol skipped from the separator. An additional gas chromatograph with a Porapack QS column connected to a TCD was then used to analyze the composition of the liquid product, after addition of a calibrated amount of acetonitrile as an internal standard. C balances with errors below 10% of the consumed carbon were always found. Adiabatic Reactor. The equipment is shown in Figure 2. Hydrogen and carbon monoxide from cylinders were mixed in the desired ratio and compressed up to the process pressure of 18 MPa. The makeup maximum flow rate was 1.5 Nm3/h. The makeup and recycle streams were mixed; a liquid stream could also be added using pump P1 before entering the preheating section where heating was provided by a 3 kW copperlined heat exchanger, E1. Reactor R1 is a 26 mm i.d. × 1100 mm Sanicro 28 Cu-lined tube. A total of 100 g of catalyst pellets was loaded inside the reactor. To allow for adiabatic operation, the reactor was placed inside a 100-mm-thick annulus made of a strongly insulating material (calcium hydrosilicate). Reaction temperatures were measured by a thermocouple sliding inside a sheath placed into the catalyst bed (TI-16). Five thermocouples were placed on the outer wall of the reactor (TI-4, TI-6, TI-8, TI-10, and TI-12) and of the insulating annulus (TI-5, TI7, TI-9, TI-11, and TI-13) as well. Five temperature controllers (TIC-5, TIC-7, TIC-9, TIC-11, and TIC-13) kept a null temperature difference between the reactor and the annulus outer wall by controlling five heating coils (RE-02 to RE-06; 0.25 kW each). Heat losses from the ending parts of the reactor were avoided by imposing a null difference along the axial direction between the temperatures of the tube wall and its ending sections (TI-19 and TI-20) and in correspondence with the catalyst bed (TI-4 and TI-12) through two heating coils (RE-01 and RE-07; 0.5 kW each). Reactor effluents were cooled to room temperature by circulating methanol in exchanger E2. Liquid product (typically 200 g/h) was collected in pot V1 and flashed to few bars in V2 where light gases were released and vented after passing through flowmeter FI-4. Uncondensed gas from V1 passed through a packed column, C1, for optional CO2 removal by methanol washing. CO2-rich methanol was collected in V4; CO2 was flashed and sent to vent by passing through flowmeter FI-5, while methanol was cooled in exchanger E3 and returned to C1 by pump P2. CO2-lean gas was recycled to R1 by recycle compressor K1 (maximum throughput 7.5 Nm3/h). A purge stream was maintained in order to avoid inerts buildup in the synthesis loop and vented after passing through flowmeter FI-3. Main gaseous streams were analyzed on-line by GC, while liquid products withdrawn from pot V2 were analyzed off-line by capillary GC. A PC equipped with the software package FIX DMACS version 2.0 (Intellution Co.) performed data collection and allowed for the remote control of the unit. The whole plant was controlled by a VF-1 IDEC PLC (Izumi) which carried out all the necessary emergency procedures in case of runaway. Design of the Experiments Table 1 reports the initial set of experiments which was designed for the kinetic study over catalyst powders. The operating field was centered around “stan-

Figure 1. Schematic of the testing unit used for the kinetic study.

3898 Ind. Eng. Chem. Res., Vol. 37, No. 10, 1998

Figure 2. Pilot-scale adiabatic unit.


3900 Ind. Eng. Chem. Res., Vol. 37, No. 10, 1998 Table 1. Design of the Experiments: Experimental Conditions Explored for the Kinetic Analysis temperature, K pressure, MPa GHSV, L(NTP)/kg of catalyst/h feed H2/CO, v/v feed CO2, % by volume

std condition

set 1 range

set 2 range

level

673 18 20 000 1.0 0.0

633-693 10-18 10 000-30 000 0.7-2.0 0.0-5.0

643-693 15-18 40 000-70 000 1.0 0.0

6 4 6 3 4

Table 2. Design of the Experiments: Additional Experimental Conditions Explored for Model Validation temperature, K pressure, MPa GHSV, L(NTP)/kg of catalyst/h feed H2/CO, v/v feed CO2, % by volume

set 3 range

level

623-703 13-23 15 000-25 000 0.7 1.0

3 2 3 1 1

dard” values for temperature, pressure, gas hourly space velocity, H2/CO feed ratio, and CO2 feed content; these were repetitively investigated during the activity tests in order to verify the catalyst stability. The ranges of the operating variables are reported in Table 1 (set 1). The maximum temperature was kept below 703 K in order to prevent catalyst sintering. In the course of the experimental research and concurrent model analysis, it turned out that extensions of the operating field toward higher space velocities and a more thorough investigation at the lower temperatures were necessary for a more reliable evaluation of the kinetics of methanol synthesis (approaching chemical equilibrium in most of the initial sets of experiments) and, as a consequence, of all the steps of the chain growth process. Additional experiments were thus designed and performed at the conditions listed in Table 1 (set 2). The overall number of levels investigated for the kinetic study is also reported in Table 1 for each operating variable; the catalyst was tested for a total period of 1000 h on stream. Finally, further kinetic runs were carried out in correspondence with “boundary” values of temperature and pressure (set 3 in Table 2); the results were not included among the data used for the kinetic analysis but were later compared with the model predictions in order to verify the model robustness to small extrapolations. Concerning the adiabatic unit, three series of experiments were performed including (i) methanol synthesis tests, (ii) higher alcohol synthesis tests with addition of methanol to the syn-gas feed, and (iii) higher alcohol synthesis tests with addition of methanol + ethanol + propanol to the feed stream. An outlet temperature of 430 °C was kept as the upper limit in operating the reactor. Among the other operating parameters, the H2/ CO makeup ratio was varied between 1 and 1.5 and GHSV between 40 000 and 80 000 NL/kg catalyst/h were tested. Reactor Models Reaction Network. The HAS reacting system was simplified into a selected number of species and pseudospecies, namely, CO, H2, CO2, H2O, methanol, ethanol, 1-propanol, isobutyl alcohol, C4+ higher alcohols (including C4 oxygenates other than isobutyl alcohol and the oxygenates with more than four carbon atoms),

methane, and C2+ hydrocarbons. The formation of the reaction products was assumed to follow the stoichiometries:

CO + 2H2 T CH3OH

(r1)

CO + H2O T CO2 + H2

(r2)

2CH3OH f C2H5OH + H2O

(r3)

CH3OH + C2H5OH f C3H7OH + H2O

(r4)

CH3OH + C3H7OH f i-C4H9OH + H2O

(r5)

(NHA - 3)CH3OH + C3H7OH f HAC4+ + (NHA - 3)H2O (r6) 2CH3OH f DME + H2O

(r7)

CO + 3H2 f CH4 + H2O

(r8)

NHYDCO + 2NHYDH2 f HYDC2+ + NHYDH2O (r9) where NHA stands for the average carbon atom number of the pseudocomponent HAC4+ (higher alcohols with four or more carbon atoms excluding isobutyl alcohol) and NHYD stands for the average carbon atom number of the pseudo-component HYDC2+ (hydrocarbons with two or more carbon atoms). A stepwise chain growth was assumed for the formation of alcohols, and methanol was formally viewed as the recurrent C1 reactant. Thus, the present chain growth process accounts for the synthesis of ethanol, 1-propanol, isobutyl alcohol, and higher oxygenates (the latter as byproduct of 1-propanol). A simplified stoichiometry was assumed for the formation of higher hydrocarbons. Methanol synthesis and the water-gas shift reaction were treated as equilibrium-constrained reactions (Tronconi et al., 1990). It is worth noticing that the dramatic reduction of the complete HAS kinetic network (Beretta et al., 1996b,c) to this simplified form is justified by the present aim, that is, the reactor design and the evaluation of the main thermal effects associated with the reaction process. The mechanistic model previously developed appeared to be redundant in this preliminary stage of the process development. Instead, preference was given to the simplified approach proposed by Tronconi et al. (1987), whose lumped kinetic model (accounting for the global formation of higher oxygenates) was herein extended; among the range of oxygenates, the present lumped kinetic model was designed to account for the production and distribution of those species which may have a real interest in a HAS process loop, both as valuable products (methanol and isobutyl alcohol) and


as recycle streams (methanol, ethanol, and propanol) (Sanfilippo et al., 1998). Kinetic Expressions. A reversible kinetic expression for methanol synthesis, accounting for the nonideal behavior of the gas mixture, was used:

r1 ) kMeOH(PCOPH22 - PMeOH/KMeOH)

(k1)

where KMeOH ) exp(-∆G°MeOH/R/T)/Kγ,MeOH ∆G°MeOH ) -24306 + 58.57T (cal/mol) and

-5

Kγ,MeOH ) 1 - P(6.713 × 10 ) exp(1.7308 × 103/T) R ) 1.987 (cal/mol/K)

The kinetics of the water-gas shift reaction were also described by an equilibrium-constrained expression; in this case nonideal behavior of the gas mixture was estimated to be negligible.

r2 ) kSHIFT(PCOPH2O - PCO2PH2/KSHIFT) where

methane and C2+ hydrocarbons formation. Tronconi et al. (1987) had already successfully used similar kinetics to describe the rate of methanation over a K-doped zinc-chromite catalyst. Indeed, as illustrated in the following, the whole bulk of experimental data could be fitted with sufficient accuracy by assuming that the rate of hydrocarbon formation was controlled by hydrogen rather than CO content.

methane synthesis r8 ) kCH4PH2

(k8)

HYDC2+ formation r9 ) kHYDPH2

(k9)

Kinetic Study: Reactor Model. The following variables were introduced: Yi ) Fi/F0. Fi (mol/h) stands for the mole flow rate of the ith species and F0 for the total inlet mole flow rate (mol/h); in correspondence with the reactor entrance, Yi0 equals the mole fraction of the ith species in the feed stream. Assuming a plug-flow model for the reactor used in the kinetic study over catalyst powders, the model included the following differential mass balances:

(k2) dYMeOH

KSHIFT ) exp(-∆G°SHIFT/R/T)

d(WC/F0)

) r1 - 2r3 - r4 - r5 - (NHA - 3)r6 - 2r7 (mb1)

∆G°SHIFT ) -8514 + 7.71T (cal/mol) In the case of the oxygenate reaction growth steps Cn + C1, simple first-order kinetic expressions with respect to Cn were used. A dependence on the hydrogen partial pressure was also included in order to account for the real nature of the reacting species, which are aldehydes and not primary alcohols (Beretta et al., 1996b). Such a mechanistic consistency was neglected only in the case of the kinetics related to the overall formation of C4+ oxygenates (an intrinsically more empirical step). Competitive adsorption terms were included to better simulate the experimental dependences of the oxygenate distribution on contact time and CO2 feed content.

ethanol formation r3 ) kC1-C2PMeOH/PH2

(k3)

PEtOH/PH2 propanol formation r4 ) kC2-C3 (k4) (1 + KH2OPH2O) isobutyl alcohol formation r5 ) kC3-iC4 × PPrOH/PH2 (1 + KH2OPH2O + KHAPHA)2 HAC4+ formation r6 ) kC3-HA supPPrOH

) r3 - r4

(mb2)

) r4 - r5 - r6

(mb3)

d(WC/F0) dYPrOH d(WC/F0)

dYiBuOH d(WC/F0) dYHA sup d(WC/F0) dYDME d(WC/F0) dYCO2

(k5)

d(WC/F0)

(k6)

dYCH4

Dimethyl ether synthesis was described with a firstorder kinetic expression in methanol partial pressure. This was preferred to a second-order expression, which was found to be less adequate in the data fit, overestimating, in particular, the effect of total pressure on the DME production.

DME formation r7 ) kDMEPMeOH

dYEtOH

(k7)

Finally, first-order dependences on hydrogen partial pressure were introduced to account for the rate of

d(WC/F0) dYHYD d(WC/F0)

) r5

(mb4)

) r6

(mb5)

) r7

(mb6)

) r2

(mb7)

) r8

(mb8)

) r9

(mb9)

with initial conditions Yi ) Yi0, for each species; WC/F0 ) catalyst load (kg)/total mole flow rate (mol/h) ) 22.4 (L(NTP)/mol)/GHSV (L(NTP)/kg of catalyst/h). The atomic balances for carbon, oxygen, and hydrogen then give rise to the following algebraic equations:


YCO ) YCO0 - (YCO2 - YCO20) - (YMeOH - YMeOH0) 2(YEtOH - YEtOH0) - 3(YPrOH - YPrOH0) 4(YiBuOH - YiBuOH0) - NHA(YHAsup - YHAsup0) 2(YDME - YDME0) + (YCH4 - YCH40) NHYD(YHYD - YHYD0) (mb10) YH2 ) YH20 + (YCO2 - YCO20) - 2(YMeOH - YMeOH0) 4(YEtOH - YEtOH0) - 6(YPrOH - YPrOH0) 8(YiBuOH - YiBuOH0) - 2NHA(YHAsup - YHAsup0) 8(YDME - YDME0) - 3(YCH4 - YCH40) 2NHYD(YHYD - YHYD0) (mb11) YH2O ) YH2O0 - (YCO2 - YCO20) + (YEtOH - YEtOH0) + 2(YPrOH - YPrOH0) + 3(YiBuOH - YiBuOH0) + (NHA - 1) (YHAsup - YHAsup0) + (YDME - YDME0) + (YCH4 - YCH40) + NHYD (YHYD - YHYD0) (mb12) The present reactor model is not strictly isothermal; the mass balances were, in fact, coupled to the axial temperature profile, represented by a 5th degree polynomial function (specific of each activity run) fitted to the experimental values of temperature measured at 12 different reactor sections. In correspondence with any reactor section, the actual mole fraction of the ith component is evaluated as

/∑

Nspecies

xi ) Yi

Yi

i)1

Adiabatic Reactor Model. This was obtained by simply coupling the mass balances (mb1)-(mb12) to the enthalpy balance

dT d(WC/FW)

N

(-∆Hiri) ∑ i)1

) 1/cP,AV

(hb1)

with the initial condition T° ) TIN. In (hb1), FW is the total mass flow rate (kg/h); cP,AV is the average heat capacity of the gas stream (assumed for simplicity constant along the reactor length and equal to the inlet value) (kcal/kg/h); and ∆Hi (kcal/mol) is the heat of the ith reaction with rate ri (mol/kg of catalyst/h). Thermodynamic data were taken from Reid et al. (1987). Kinetic Study The parameters included in the kinetic model [(k1)(k9) + (mb1)-(mb12)] as well as their dependences on the reaction temperature were estimated by multiresponse nonlinear regression on the results of 25 experimental runs selected from sets 1 and 2, providing a uniform coverage of the experimental field. Independent replicated runs were used to estimate the experimental error variances for the model responses. The data account for the effects of temperature, pressure, gas hourly space velocity, H2/CO synthesis gas ratio, and CO2 feed content on the product yield and distribution. The variables used as model responses in

Figure 3. Experimental and calculated trends of methanol concentration with varying operating conditions. (a) GHSV ) 20 000 L(NTP)/kg of catalyst/h, P ) 1.8 MPa, H2/CO ) 1, CO2 ) 0%. (b) T ) 673 K, P ) 1.8 MPa, H2/CO ) 1, CO2 ) 0%. (c) T ) 673 K, GHSV ) 20 000 L(NTP)/kg of catalyst/h, H2/CO ) 1, CO2 ) 0%. (d) T ) 673 K, GHSV ) 20 000 L(NTP)/kg of catalyst/h, P ) 1.8 MPa, CO2 ) 0%. (e) T ) 673 K, GHSV ) 20 000 L(NTP)/kg of catalyst/h, P ) 1.8 MPa, H2/CO ) 1.

the regression procedure were the outlet percentage mole fractions of methanol, CO2, dimethyl ether, ethanol, propanol, isobutyl alcohol, methane, higher oxygenates, and higher hydrocarbons. The results are reported in Figures 3-9 which compare experimental (symbols) and calculated (curves) values of the outlet concentration (% mol) of the various reaction products. Methanol. A complete approach to chemical equilibrium was reached in correspondence with high temperatures and low space velocities (this involves part of the data reported in Figure 3a,b, as well as the data represented in Figure 3c-e which account for the effects of pressure and feed composition at high temperature and contact time). This was nicely reproduced by the model through the reversible kinetics herein assumed. The runs performed at lower temperature and higher space velocity (especially 70 000 L(NTP)/kg of catalyst/ h) showed, instead, a significant deviation of methanol production from the chemical equilibrium. These experiments were greatly informative with regard to the estimation of the kinetic constants included in (k1). DME. Its rate of formation appeared sufficiently well reproduced by a simple first-order dependence on PMeOH.


Figure 4. Experimental and calculated trends of ethanol concentration with varying operating conditions.

Data are not reported for brevity. The effects of temperature, contact time, and feed gas composition on the production of DME were strictly similar to the corresponding trends of methanol concentration. Ethanol and Propanol. In all of the experiments, the outlet concentration of ethanol was always very low, as reported in Figure 4. Also, its dependence on the operating variables appeared very weak. This is a peculiar feature of ethanol (Beretta et al., 1996b,c), related to the special role of this species (a highly reactive intermediate) in the chain-growth process. The present model does not reproduce the details of the complex reaction network which involves ethanol in higher alcohol synthesis; ethanol was simply schematized as a secondary product of methanol and a precursor of propanol. Still, the resulting model fit was satisfactory. Propanol production was more sensitive to the variation of the operating conditions. As reported in Figure 5, propanol outlet concentration increased considerably with decreasing space velocity and increasing total pressure, while it was moderately promoted by increasing reaction temperature and hydrogen feed content. No significant inhibiting effect was exerted by CO2 addition. The model accuracy in the description of such effects was especially good. Isobutyl Alcohol. Figure 6 shows the experimental and calculated trends of isobutyl alcohol concentration upon varying the operating conditions. Temperature,

Figure 5. Experimental and calculated trends of propanol concentration with varying operating conditions.

contact time, and pressure exerted the most pronounced effects. Concerning the model calculations, the inclusion of the adsorption terms for oxygenates and water in the denominator of expression (k3) was necessary for an adequate simulation of the dependence of isobutyl alcohol production on space velocity. Through the reverse water-gas shift reaction, which accounts for the conversion of CO2 to CO and water, the same term allowed us to describe accurately the inhibiting effect of CO2 feed enrichment on isobutyl alcohol production. Higher Alcohols and CO2. The overall production of higher oxygenates is reported in Figure 7 as a function of the process variables. It increased with increasing temperature, contact time, and pressure. It was almost unaffected by the H2/CO ratio for values close to unity and only weakly retarded by the addition of CO2. The model fit was, in general, satisfactory except for a certain degree of underestimation of the higher oxygenates in correspondence with the highest temperature level. The formation of oxygenates was accompanied by the production of water and, through the water-gas shift reaction, of CO2. The observed trends of CO2 concentration were accordingly similar to the corresponding trends of oxygenates production. The model adequacy was confirmed also in this case. CH4 and C2+ Hydrocarbons. A low production of hydrocarbons was found in all of the experiments.


Figure 6. Experimental and calculated trends of isobutyl alcohol concentration with varying operating conditions.

Figures 8 and 9 show the dependence of hydrocarbon production on the operating variables which were observed during the HAS runs. Such dependences were very nicely described by the model, despite the simplified nature of kinetics (k8) and (k9) herein adopted. The data and the model predictions indicated that low temperatures and pressures as well as short contact times and H2/CO ratios below 1 reduce the production of methane and higher hydrocarbons. Optimal estimates and confidence limits for the kinetic parameters are reported in Table 3. The extradiagonal component of the correlation matrix with the highest absolute value amounted to 0.83; a general absence of strong correlation between parameter estimates was, in fact, detected. Comparison with Additional Data Data analysis also involved the simulation of nine activity runs, representative of set 3 of the experimental plan. The comparison of model predictions with the measured product distributions was meant to verify the quality of small extrapolations with respect to the central experimental field where the model was fitted to data. Figure 10 shows the overall result. The errors associated with the model predictions for set 3 appeared comparable with the average error of the model fits for sets 1 and 2. The highest deviations were related to model extrapolations to a pressure of 23 MPa, where a

Figure 7. Experimental and calculated trends of the total higher oxygenate concentration with varying operating conditions.

certain degree of overestimation of hydrocarbon production was found. Model predictions (in the case of both oxygenates and hydrocarbon production) were instead more accurate in the field of low pressures (13 MPa). Simulation of an Adiabatic Unit Once the kinetic parameters were estimated, the model accounting for the enthalpy balance (hb1) was applied to the simulation of an adiabatic multistage reactor in order to verify on a preliminary basis the feasibility of an industrial unit. Special attention was paid to the expected thermal behavior of the reactor; this was investigated through a sensitivity analysis by perturbing the simulated operating conditions of the reactor (inlet temperature, total pressure, catalyst load, feed composition). An upper bound on catalyst temperature was assumed, equal to 693 K. The most remarkable result was finding that over the present Snamprogetti catalyst the overall process of exothermicity was governed by methanol synthesis. The simulations showed that the extent of heat production could be moderated by adding a proper amount of methanol to the synthesis gas feed stream. Such an amount was strictly dependent on the operating conditions. Assuming a total pressure P ) 18 MPa, inlet temperature TIN ) 370 °C, and H2/CO ) 1 in the feed


Figure 8. Experimental and calculated trends of CH4 concentration with varying operating conditions.

stream, it was estimated that the catalyst temperature profile could be maintained below the limit of 693 K for a space velocity of practical interest only if 2.5% v/v methanol was cofed to the reactor. For this case, Figure 11 shows the calculated temperature profile and the calculated profiles of methanol, isobutyl alcohol, and methane concentrations along three adiabatic layers, assuming interstage cooling to 643 K. The temperature profile of each stage exhibited a characteristic point where the gradient changed dramatically (Figure 11a). At the reactor entrance, in fact, the calculated temperature rise was sharp and 683 K was reached after about one-fourth of the whole bed length; along the rest of the catalytic bed, instead, the temperature increased only moderately (10-15 K). This could be easily explained by considering the trends of the concentrations of major products (Figure 11b,c). The formation of higher oxygenates and hydrocarbons was progressive along the reactors; the corresponding resulting heat production thus distributed throughout the entire catalytic bed. On the contrary, the production of heat accompanying methanol synthesis was concentrated within the entry region of the reactor. Here, based on the model simulation, methanol concentration grew up from the inlet to the equilibrium level; the enthalpy contribution was thus strictly related to the extent of the gap between inlet and equilibrium concentrations. After the reactor section where the thermodynamic constraint was reached, the contribution of methanol synthesis to the exothermicity dropped to zero. The residual weaker

Figure 9. Experimental and calculated trends of the total C2+ hydrocarbons concentration with varying operating conditions.

temperature increase resulted from the exothermicity of the other reactions. This was further moderated by an endothermic contribution associated with the reverse methanol synthesis due to the negative temperature dependence of methanol equilibrium. The shape of the temperature and concentration profiles seemed to rule out the risk of reactor runaway. This was an especially promising result. Fixed-bed multitubular reactors and slurry phase reactors were also considered as alternative solutions in the very beginning of the project, because they allow optimal temperature control. However, a multiple-fixed-bed adiabatic reactor with intermediate cooling was preferred for the following main reasons: (i) it can be designed to accommodate larger alcohol productions than those obtainable in a multitubular reactor since the catalyst inventory per single unit can be larger; (ii) reaction rates are not diffusion-limited as in a slurry reactor, thus resulting in a more efficient utilization of the reactor volume. HAS Runs in the Adiabatic Pilot Reactor The adiabatic unit was first tested at low inlet temperature (≈543 K), that is, under conditions favorable to the selective formation of methanol. The reactor showed an unstable behavior, as the axial temperature profile typically presented a sharp T increase to 693 K near the outlet of the reactor. As a consequence, even

3906 Ind. Eng. Chem. Res., Vol. 37, No. 10, 1998 Table 3. Vector of the Kinetic Parameters and Optimal Estimatesa x(1) x(2) x(3) x(4) x(5) x(6) x(7) x(8) x(9) x(10) x(11) x(12) x(13) x(14) x(15) x(16) x(17) x(18) x(19) x(20)

k0MeOH 0 kC1-C2 0 kC2-C3 0 kC3-iC4 0 kC3-HAsup 0 kCH4 k0HYD 0 KH 2O 0 K HA k0DME Eatt,MeOH/R Eatt,C1-C2/R Eatt,C2-C3/R Eatt,C3-iC4/R Eatt,C3-HAsup/R Eatt,CH4/R Eatt,HYD/R -(∆Hads,H2O/R) -(∆Hads,HA/R) Eatt,DME/R

(8.077 ( 1.683) × 10-4 (4.185 ( 0.184) × 101 (1.122 ( 0.207) × 104 (1.038 ( 0.229) × 104 (2.834 ( 0.296) × 101 (1.348 ( 0.159) × 10-2 (1.231 ( 0.109) × 10-2 (1.556 ( 0.607) (1.303 ( 0.263) (2.146 ( 0.259) × 10-1 (1.626 ( 0.088) × 101 (1.412 ( 0.062) × 101 (2.974 ( 0.595) (7.198 ( 2.839) (2.114 ( 3.259) × 10-3 (9.068 ( 1.062) (1.422 ( 0.138) × 101 (2.513 ( 0.124) × 101 (2.109 ( 3.219) × 10-1 (6.620 ( 0.360)

a The following form was used: k ) ko exp[-E i att,i/R(1000/T i 1000/To)] with To ) 673 K. koi dimensions depend on the kinetic expression assumed for the corresponding ri which is expressed as mol/kg of catalyst/h. 95% confidence limits were estimated by assuming the following values of the response variances: σMeOH2 ) 1.4 × 10-1, σCO22 ) 2.25 × 10-1, σEtOH2 ) 5.0 × 10-5, σPrOH2 ) 5.0 × 10-5, σiBuOH2 ) 1.0 × 10-3, σHA2 ) 1.0 × 10-2, σCH42 ) 2.2 × 10-3, σHYD2 ) 1.0 × 10-3, σDME2 ) 4.0 × 10-3. kSHIFT ) 4.21 × exp[-3.491(1000/T - 1000/To)] was estimated independently from the other kinetic parameters.

Figure 11. Simulation of a three-stage adiabatic plug-flow reactor. P ) 18 MPa. In each stage: Tinlet ) 643 K, Toutlet ) 693 K. First-stage feed stream: H2/CO ) 1, methanol ) 2.5% v/v. R1: GHSV ) 100 000 L(NTP)/kg of catalyst/h. R2: GHSV ) 55 500 L(NTP)/kg of catalyst/h. R3: GHSV ) 55 500 L(NTP)/kg of catalyst/h. In Figure 12b the solid lines represent the calculated axial profiles of the equilibrium methanol concentration.

Figure 10. Simulation of the experimental data of set 3 and comparison with the model fit to sets 1 and 2. Symbols represent the experimental and calculated values of the percentage mole fractions of methanol, ethanol, propanol, isobutyl alcohol, total higher alcohols, DME, methane, and C2+ hydrocarbons.

small perturbations in the feed flow rate could result in dramatic changes of the outlet reactor temperature and methanol concentration. On the contrary, the operation of the adiabatic reactor under higher alcohol synthesis conditions was extremely stable. In agreement with the simulations previously discussed, it was confirmed that smooth S-shaped temperature profiles could be obtained by cofeeding

methanol to the synthesis gas feed stream. The proper amount of methanol feed content (between 1 and 5% v/v) depended on the imposed operating conditions. Figures 12 and 13 show in symbols the measured temperature profiles during a series of runs wherein GHSV, inlet reactor temperature, methanol (and ethanol + propanol) feed content, and H2/CO synthesis gas ratio had been changed. The data were fitted by the same adiabatic reactor model used for the simulations discussed in the previous section. In view of the purposes of the present work, no effort was made at this stage to account for such effects as intraporous diffusional limitations and axial dispersion of heat and mass, which were expectedly different at the laboratory and pilot scale. This resulted in some adjustment of the parameter estimates as compared to their original


Figure 12. Experimental and calculated axial temperature profiles of the adiabatic unit. P ) 18 MPa. (a) Run 1: Tinlet ) 601 K, GHSV ) 80 000 L(NTP)/kg of catalyst/h. Feed: inlet H2/CO ) 1.65, CO2 ) 0.64%, methanol ) 2.1% by volume (b) Run 2: Tinlet ) 614 K, GHSV ) 80 000 L(NTP)/kg of catalyst/h. Feed: inlet H2/ CO ) 1.16, CO2 ) 1.46%, methanol ) 2.3% by volume. (c) Run 7: Tinlet ) 634 K, GHSV ) 65 000 L(NTP)/kg of catalyst/h. Feed: inlet H2/CO ) 1.3, CO2 ) 1.98%, methanol ) 2.9% by volume.

values given in Table 3. The most significant change was the decrease of the activation energy for methanol synthesis from 32 to about 22 kcal/mol; this was introduced after fitting the model to a number of especially informative runs, performed at high GHSV and low inlet temperature. The solid lines in Figures 12 and 13 represent the model predictions. Experimental and calculated data demonstrated that the reactor could be maintained under predetermined and controlled T conditions by simply manipulating methanol recycle at varying flow rate, temperature, and composition of the feed stream. As the overall production of heat was proportional to the net amount of methanol produced along the catalyst bed, increasing amounts of methanol were cofed at increasing inlet temperature (from parts a to b to c in Figure 12) in order to keep the outlet temperature at about 693 K and the end portion of the temperature profile almost flat. At increasing GHSV (from parts c to b to a in Figure 13),

Figure 13. Experimental and calculated axial temperature profiles of the adiabatic unit. P ) 18 MPa. (a) Run 15: Tinlet ) 618 K, GHSV ) 80 000 L(NTP)/kg of catalyst/h. Feed: inlet H2/ CO ) 1.48, CO2 ) 2.53%, methanol ) 2.24%, ethanol ) 0.027%, propanol ) 0.039% by volume. (b) Run 4: Tinlet ) 630 K, GHSV ) 65 000 L(NTP)/kg of catalyst/h. Feed: inlet H2/CO ) 1.25, CO2 ) 2.02%, methanol ) 2.84% by volume. (c) Run 14: Tinlet ) 640 K, GHSV ) 41 000 L(NTP)/kg of catalyst/h. Feed: inlet H2/CO ) 1.46, CO2 ) 2.23%, methanol ) 4.3%, ethanol ) 0.052%, propanol ) 0.076% by volume.

instead, the inlet methanol content was decreased in order to increase the rate of heat production and guarantee a constant outlet temperature with reduction of the contact time. A cofeed of low amounts of higher alcohols (0.010.08% v/v), namely, ethanol and propanol, did not produce any significant change in the thermal behavior of the adiabatic reactor. However, an increase of isobutyl alcohol production was predicted and observed. Methane and higher hydrocarbon concentrations were always low in the product mixture and in line with the kinetic data previously discussed. The general good agreement between experimental and calculated T profiles showed a remarkable approach of the adiabatic reactor to the ideal operation; the observed T increase (∆T) was, in fact, always between 95% and 100% of the ideal adiabatic ∆T.


Conclusions The development of a high-performance higher alcohol synthesis catalyst has stimulated a process design research. This was initiated with a broad kinetic investigation. Activity runs over small catalyst particles were interpreted and quantitatively accounted for by a simplified model, wherein the detailed description of the reaction mechanism was partially neglected in favor of an “easy-to-handle” mathematical tool, more suitable to engineering applications. An adiabatic pilot-scale reactor was then realized. It was demonstrated that methanol synthesis controlled the overall thermicity of the process. Therefore, due to the thermodynamic constraints prevailing on methanol synthesis, risks of runaway can be drastically reduced by adding a suitable recycle of methanol into the synthesis gas feed stream. Also, the possibility of uncontrolled heat production due to the formation of hydrocarbons which has been suggested in the literature could be ruled out over the present zinc-chromite catalyst given the low productivity of methane and higher hydrocarbons. Literature Cited Antonelli, G.; Cornaro, U. Procedimento per produrre miscele di metanolo e alcoli superiori. Italian Patent MI95 A 002100, 1995. Beretta, A.; Sun, Q.; Herman, R. G.; Klier, K. Production of Methanol and Isobutyl Alcohol Mixtures over Double-Bed Cesium-Promoted Cu/ZnO/Cr2O3 and ZnO/Cr2O3 Catalysts. Ind. Eng. Chem. Res. 1996a, 35, 1534-1542. Beretta, A.; Tronconi, E.; Forzatti, P.; Pasquon, I.; Micheli, E.; Tagliabue, L.; Antonelli, G. B. Development of a Mechanistic Kinetic Model of the Higher Alcohol Synthesis over a Cs-Doped Zn/Cr/O Catalyst. 1. Model Derivation and Data Fitting. Ind. Eng. Chem. Res. 1996b, 35, 2144-2153. Beretta, A.; Lietti, L.; Tronconi, E.; Forzatti, P.; Pasquon, I. Development of a Mechanistic Kinetic Model of the Higher Alcohol Synthesis over a Cs-Doped Zn/Cr/O Catalyst. 2. Analysis of Chemical Enrichment Experiments. Ind. Eng. Chem. Res. 1996c, 35, 2144-2153. Forzatti, P.; Tronconi, E.; Pasquon, I. Higher Alcohol Synthesis. Catal. Rev.-Sci. Eng. 1991, 33, 109-168. Hindermann, J. P.; Hutchings, G. J.; Kiennemann, A. Mechanistic Aspects of the Formation of Hydrocarbons and Alcohols from CO Hydrogenation. Catal. Rev.-Sci. Eng. 1993, 35, 1-127.

Nunan, J. G.; Bogdan, C. E.; Klier, K.; Smith, K. J.; Young, C.W.; Herman, R. G. Methanol and C2 Oxygenate Synthesis over Cesium Doped Cu/ZnO and Cu/ZnO/Al2O3 Catalysts: A Study of Selectivity and 13C Incorporation Patterns. J. Catal. 1988, 113, 410-433. Ohno, T.; Yoshimoto, M.; Asselineau, L.; Courty, T.; Travers, P. AIChE Spring National Meeting, New Orleans, LA, 1986. Paggini, A.; Sanfilippo, D.; Pecci, G.; Dybkjaer, I. Implementation of the methanol plus higher alcohols process by Snamprogetti, Enichem, Haldor Topsøe a/s “Mas Technology”. VII International Symposium on Alcohol Fuels, Paris, Oct 20-23, 1986. Reid, R. C.; Prausnitz, J. M.; Poling, B. E. The Properties of Gases and Liquids, 4th ed.; McGraw-Hill: New York, 1987. Sanfilippo, D.; Micheli, E.; Miracca, I.; Tagliabue, L. Oxygenated Synfuels from Natural Gas. Pet. Technol. Q. Spring 1998, 8795. Smith, K. J.; Young, C.-W.; Herman, R. G.; Klier, K. Development of a Kinetic Model for Alcohol Synthesis over a CesiumPromoted Cu/ZnO Catalyst. Ind. Eng. Chem. Res. 1991, 30, 6171. Stiles, A. B.; Chen, F.; Harrison, J. B.; Hu, X.; Storm, D. A.; Yang, H. X. Catalytic Conversion of Synthesis Gas to Methanol and Other Oxygenated Products. Ind. Eng. Chem. Res. 1991, 30, 811-821. Tronconi, E.; Ferlazzo, N.; Forzatti, P.; Pasquon, I. Synthesis of Alcohols from Carbon Oxides and Hydrogen. 4. Lumped Kinetics for the Higher Alcohol Synthesis over a Zn-Cr-K Oxide Catalyst. Ind. Eng. Chem. Res. 1987, 26, 2122-2129. Tronconi, E.; Forzatti, P.; Pasquon, I. I. An Investigation of the Thermodynamic constraints in Higher Alcohol Synthesis over Cs-Promoted ZnCr-Oxide Catalyst. J. Catal. 1990, 124, 376390. Tronconi, E.; Lietti, L.; Groppi, G.; Forzatti, P.; Pasquon, I. Mechanistic Kinetic Treatment of the Chain Growth Process in Higher Alcohol Synthesis over a Cs-Promoted Zn-Cr-O Catalyst. J. Catal. 1992, 135, 99-114. Underwood, R. P.; Schaub, E. Synthesis Gas Conversion to Isabutanol: a Key Step in a C1 Route to MTBE. 10th International Symposium on Alcohol Fuels, Colorado Springs, CO, Nov 7-11, 1993. Zhou, P. Summary of the Higher Alcohol Synthesis Workshop; U.S. Department of Energy: Pittsburgh, PA, Feb 1994.

Received for review October 21, 1997 Revised manuscript received May 28, 1998 Accepted July 9, 1998 IE9707331

Development of a Process for Higher Alcohol Production via Synthesis

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