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Znd. Eng. Chem. Res. 1993,32, 1610-1621
Deactivation of Methanol Synthesis Catalysts George W.Roberts,’ Dennis M. Brown, Thomas H. Hsiung, and J o h n J. Lewnard Air Products and Chemicals, Znc., 7201 Hamilton Boulevard, Allentown, Pennsylvania 18195-1501
A novel methanol synthesis process, the liquid-phase methanol (LPMEOH) process, has been developed and scaled up t o a nominal 380 kg/h (10 ton/day) pilot plant. The process is based on a gas-sparged slurry reactor instead of a conventional, fixed-bed reactor. The use of slurry reactors, which are essentially gradientless, greatly facilitated the interpretation and quantification of catalyst deactivation phenomena. With a poison-free, CO-rich feedstream, the rate of deactivation of the Cu/ZnO catalyst increased rapidly with temperature. At constant temperature, in the absence of poisons, the decline with time in the rate constant for methanol synthesis correlated with the loss of BET surface area. Iron carbonyl, nickel carbonyl, and carbonyl sulfide are severe and highly specific poisons for methanol-synthesis catalyst. There was a linear relationship between the catalyst activity loss and the concentration of metal or sulfur on the catalyst. Introduction The actual and potential emergence of several major new applications for methanol has led to renewed interest in the technology for producing this chemical. Substantial quantities of methanol are currently being used for the production of methyl tert-butyl ether (MTBE), a rapidlygrowing octane enhancer for gasoline that also serves to reduce carbon monoxide and hydrocarbon emissions from automobiles. Potential uses for methanol include as a transportation fuel, either neat or blended with gasoline, as a fuel for gas turbine engines for generation of electric power, as a fuel for stationary boilers and as a feedstock for the methanol-to-gasoline (MTG) process. The potential use of methanol as a fuel is being driven, in part, by its low-emissions characteristics. One of the newest methanol production scenarios involves coal gasification combined cycle (CGCC) power plants. As shown in Figure la, CGCC plants consist of three major sections: coal gasification, heat recovery/ impurity removal, and power generation via gas and steam turbines. A disadvantage of this configuration involves load following, i.e., changing the output of the plant in response to variations in the demand for electricity. For the flowsheet of Figure la, load following can be done only by varying the feed rates of 02,H20, and coal to the coal gasifier. However, since the first two sections of the plant are very capital intensive, reducing their output results in poor utilization of capital, raising the unit cost of synthesis gas and electric power. Figure l b shows a modification of the basic CGCC flowsheet. A methanol-synthesis section has been inserted between the heat recovery/gas cleanup and turbine sections, thus permitting the coproduction of methanol and electric power. In the process of Figure lb, the coal gasifier is always run at design capacity and the clean synthesis gas (syngas)is distributed between the methanol-synthesis section and the turbine section according to the power demand (Brown et al., 1988; Moore et al., 1989). During periods of low power demand, most or all of the syngas is fed to methanol synthesis. The methanol produced goes to storage and the unconverted syngas is fed to the gas turbines. During periods of peak power demand, all of the synthesis gas is fed to the turbine section and methanol can be taken from storage and burned as an auxiliary fuel,
* To whom correspondence should be addressed. Current address: Department of Chemical Engineering, North Carolina State University, Box 7905, Raleigh, NC 27695-7905.
if necessary. At the design power load, the split of clean syngas between the methanol synthesis and turbine sections is determined by the desired annual production rate of merchant methanol, which also determines the sizing of the gasifier, the heat recovery/gas cleanup, and the methanol-synthesis sections. The integration of methanol synthesis into a CGCC plant imposes several important requirements on the methanol process of Figure lb. First, it should produce high conversions to methanol on a “once-through” basis, i.e., without recycle of unconverted reactants. The extent to which power output can be reduced without reducing the output of the coal gasifier is largely determined by the fraction of the energy content in the syngas entering the methanol-synthesis section that can be converted to methanol. Second, the process should be compatible with modern, thermally-efficient coal gasifiersof the entrainedflow, slagging type. Such gasifiers produce synthesis gas with HdCO ratios in the range of 0.5-1, far removed from ratios exceeding 2.0 that are representative of the fresh feed to conventional methanol-synthesis reactors. The methanol process of Figure l b should be able to accept a CO-rich synthesis gas directly from the gas cleanup/heat recovery section, preferably without the expensive watergas-shift step that is commonly used in vapor-phaseplants prior to the methanol-synthesis section. The liquid-phase methanol process is a major departure from traditional process designs based on fixed-bed catalytic reactors. It addresses the special CGCC requirements described above and also has the potential to provide significant advantages in the conventional context of “stand-alone” methanol plants. The formation of methanol from hydrogen and carbon oxides is highly exothermic. At temperatures and pressures that are characteristic of modern methanol plants, the equilibrium conversion of the limiting reactant, oxides of carbon, is substantially less than 100% and it declines sharply with temperature. One of the most difficult design problems is removing the heat of reaction rapidly enough to maintain the close temperature control that is required for high reactant conversion and long catalyst life. In conventional, fixed-catalyst-bed processes, temperature control is partially achieved by employing a dilute, Hz-rich feedstock to limit the per-pass production of methanol and, thereby, the reaction exotherm. For example, the fresh feed to a typical, fixed-bed methanol-synthesis-reactor system has a stoichiometric ratio, defined as Hz/(CO + 1.5C02), of about 2.5. At this ratio, the syngas contains 25% more Hz than is required to react with the carbon oxides. One part
0888-5885/93/2632-1610$04.00/0 0 1993 American Chemical Society
Ind. Eng. Chem. Res., Vol. 32, No. 8, 1993 1611
H
t
S, TARS, ETC.
b
STEAM
COAL 0 2
n
COAL -- GASlFlCATlON
-
e
H
HEAT RECOVERY AND IMPURITY REMOVAL
TAM
-
S, TARS, ETC.
MeOH STORAGE
POWER
GAS TURBINES
SYNTHESIS
L
MERCHANT MeOH
i
Figure 1. Electric power by coal gasification combined cycle (CGCC) process. (a) Basic flowsheet. (b) Coproduction of power and methanol.
of this fresh feed is combined with about 5 parts of recycled effluent from the reactor system, after removal of water and methanol, to form a total reactor feed that has a high stoichiometric ratio, on the order of 10, and is very dilute in the carbon oxides, on the order of 6 mol % total (Islam and Earl, 1990). Although the excess Hz helps to shift the reaction equilibrium toward methanol, the low concentration of carbon oxides limits the amount of methanol that can be produced. The LPMEOH process is based on the use of a slurry reactor. The methanol-synthesis catalyst is suspended in an inert hydrocarbon liquid through which the feed gas is sparged. The inert liquid absorbs the reaction heat very effectively from the small (1-10pm) catalyst particles, and the backmixed nature of the slurry reactor provides avery close approach to isothermal operation. The slurry reactor can be operated with much higher concentrations of limiting reactant in the feed, leading to much higher concentrations of product in the effluent. A single slurry reactor has been run with up to 15 mol % methanol in the reactor effluent. Typically, the concentration of methanol in the effluent from a fixed-bed, vapor-phase process is less than about 6 mol % (Supp, 1973). The LPMEOH process for methanol synthesis has been compared to conventional, fixed-bed processes in more detail in earlier papers (Brown, 1984; Roberts et al., 1985). This paper is concerned with the deactivation of methanol-synthesis catalyst in a slurry reactor. A good deal of information is available concerning catalyst deactivation in conventional, fixed-bed processes (Thomas, 1970; Supp, 1973; Stiles, 1977; Satterfield, 1980). It is well-known that methanol-synthesis catalysts can be poisoned by various species including metal carbonyls and compounds containing sulfur and chlorine, and that these catalysts can also deactivate via a thermally-driven sintering mechanism. Nevertheless, there were several new aspects to the present study including (1)the use of
a concentrated, CO-rich feed gas, (2) the presence of a hydrocarbon liquid in contact with the catalyst, and (3) the need for amore detailed understanding of the poisoning behavior of species peculiar to CO-rich gas, in particular iron carbonyl,nickel carbonyl, and carbonylsulfide. These volatile compounds have a greatly-increased tendency to form at the high carbon monoxide partial pressures that are characteristic of the CGCC application. Finally, the slurry reactors used in this research were highly backmixed. Temperature and concentration were essentially uniform throughout, so that all catalyst particles were exposed to the same process conditions a t any point in time. The use of these “gradientless” reactors made it possible to define very precisely the effect of temperature and the effect of poisons on the catalyst deactivation rate. The data obtained from these reactors appear to contradict some of the conclusions from previous studies on the deactivation of methanol-synthesis catalysts. Experimental Equipment and Procedures
The catalyst deactivation studies described in this paper were carried out in laboratory reactors and in a nominal 380 kg (CH30H)/h (10 ton/day) pilot-plant reactor located in La Porte, TX. This slurry bubble column reactor had an inner diameter of 57 cm and a filled length of between 549 and 610 cm. The laboratory experiments were done in 300-cm3,gas-sparged, mechanically-agitated autoclaves. One series of laboratory studies was carried out in a 1.2cm inner-diameter tubular reactor operating as a conventional, vapor-phase, fixed catalyst bed. The details of the design and operation of these reactor systems have been reported elsewhere (Brown and Greene, 1984; Air Products, 1986; Lewnard et al., 1986). Most of the present studies were carried out with a feed gas composition referred to as “CO-rich gas”. A limited number of experimentswere carried out with a composition
1612 Ind. Eng. Chem. Res., Vol. 32, No. 8, 1993 Table I. Feed Gas Compositions for Catalyst Aging Studies CO-rich gas balanced gas component (mol ’3%)
Hz
CO
coz
35 51 13
CHI
Nz
stoichiometric ratioa Defined as H2/(CO
1 0.50
55 10 5 2 19 2.08
+ 1.5co2).
designated as “balanced gas”. These compositions are shown in Table I. The stoichiometric ratio of CO-rich gas is representativeof the gas produced by modem, thermallyefficient coal gasifiers such as the entrained-flow,slagging gasifiers of Texaco, Shell, and DOW,and the composition is typical of the feed to a “once-through” methanol unit in a CGCC power plant containing a Texaco gasifier. The composition of balanced gas is representative of the fresh feed to a stand-alone, “all-methanol” plant based on the slurry reactor concept, with coal as the feedstock and with water-gas shift prior to the methanol-synthesis loop. Balanced gas has a relatively high N2 content because N2, an impurity in the 0 2 that is fed to the coal gasifier, builds up in the recycle loop in an “all-methanol” plant. The composition of balanced gas is slightly on the hydrogenrich side of stoichiometric. However, it has a lower stoichiometric ratio than a typical feed to a conventional, vapor-phase reactor. Both CO-rich and balanced gas have concentrations of H2 and CO that are high relative to conventional, vaporphase feedstreams, and as discussed above, both feeds have the potential to produce very high concentrations of methanol in the outlet stream. In the laboratory experiments, a series of adsorption columns was used to remove known catalyst poisons and ensure a contaminant-free feed gas (Golden et al., 1991). The effluent from the adsorption system was monitored by gas chromatography. Except as specificallymentioned, the feed concentrationsof known poisons such as hydrogen sulfide, carbonyl sulfide, chlorine compounds, iron carbonyl, and nickel carbonyl were below the normal limits of chromatographic detection, Le., 15, 30, 10, 0.3, and 1 ppb by volume, respectively. Since the metal carbonyls were of specific concern in this research, occasional spot checks were made on the concentrations of these species in the feed gases using a more laborious ‘bubbler” technique (Air Products, 1987). The results showed that the adsorption columns removed iron and nickel to a concentration of less than 1 ppb(v). The Hz and CO that were fed to the La Porte pilot plant were obtained as separate components from Air Products and Chemicals plant at La Porte, which is adjacent to the pilot plant. These species were originally manufactured by steam reforming of natural gas. Three different low-pressure methanol-synthesis catalysts of the Cu/ZnO/Al203 type were used during this study. They were obtained from commercial manufacturers under agreements that preclude disclosure of compositional or structural information beyond what is given in this paper and in other published documents (Karwacki et al., 1984). Several different hydrocarbon liquids were used as slurry media; all were paraffinic or paraffinic/naphthenic oils with a boiling range from about 525 to 825 K. Analysis of used catalyst samples for iron and nickel was done by atomic absorption spectroscopy. Catalyst deactivation was followed by measuring the rate of reaction as a function of time for a fixed set of
operating conditions. The overall rate of methanol , calculated formation per unit weight of catalyst, r ~was directly from the measured inlet and outlet concentrations and flow rates, and from the known weight of catalyst in the reactor. This parameter is frequently referred to as the catalyst productivity. For qualitative comparisons, the decline in the catalyst productivity with time was compared from one experiment to another. For quantitative comparisons, and for the purpose of developing a mathematical model of catalyst deactivation, a rate constant was calculated from r~ using the rate equation (Lewnard and Hsiung, 1987;Weimer et al., 1987)
In performing this calculation, it was assumed that the reactor was perfectly backmixed, i.e., that the reactant and product partial pressures were uniform throughout the reactor and equal to those in the exit stream. This assumption is not rigorously valid for the pilot-plant reactor (Brown, 1984), but it does not distort the conclusions of this study. This rate equation is much less sophisticated than more complex kinetic models that account for the effects of C02, inhibition by product methanol, and inhibition by H2 and CO (Stiles, 1977; Supp, 1990). However, in the present study, eq 1 was used with a constant feed composition (CO-rich gas), with a constant pressure (5.27 MPa), and over a relatively limited range of temperature and conversion. In this context, eq 1 proved to be an adequate representation of the reaction kinetics. Results Qualitative Studies. There were two very fundamental questions related to catalyst deactivation that impacted basic process feasibility: if (1)the presence of a liquid phase or (2) the use of a concentrated, CO-rich feed gas had caused rapid catalyst deactivation, it would not have been reasonable to undertake large-scale process development. These two questions were addressed early in the research program, prior to the evolution of complete protocols for gas analysis, catalyst analysis, and data evaluation. The use of concentrated, CO-rich gas was of particular concern initially. On the basis of conventional wisdom in the methanol industry, CO-rich gas was expected to accelerate the rate of catalyst deactivation because of the potential for carbon formation and/or high catalyst temperatures. The use of a slurry reactor, or any other highly backmixed reactor, could have aggravated any effect associated with carbon formation since the outlet gas, to which all of the catalyst is exposed, is even more CO-rich than the inlet gas. At “normal” conversions of CO-rich gas, the HdCO ratio and the stoichiometric ratio of the outlet gas from the slurry are typically about 0.30 and 0.25, respectively. 1. Gas-Phase versus Liquid-Phase Operation. In order to explore the influence of gas-phase versus liquidphase operation, catalyst life tests were conducted in a gas-phase tubular reactor and in a stirred autoclave using the same catalyst under similar operating conditions. The nominal temperature (523 K), the total pressure (5.27 MPa (750psig)),and the space velocity (5,000 L/(kg(catalyst).h) were the same in both experiments. The catalyst was in powdered form, and balanced gas was used in both experiments. Balanced gas was chosen for this test partly because its stoichiometry is closer to that of a conventional feed gas than the stoichiometry of CO-rich gas, thus
Ind. Eng. Chem. Res., Vol. 32, No. 8,1993 1613
I”. - 0
Llquld Phase
i ? -
.
I
- - + - - Gas Phase
4i
h
5
L
O! 0
500
1000
0.8
1500
Time, hr.
Figure 2. Catalyst activity maintenance in liquid phase and gas phase: 5.27 MPa, 523 K, 5000 L/(kg(catalyt).h), balanced gas, 15 wt % slurry. 20,
.
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Autoclave 5,000 Vkg-hr 15 wt% slurry
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