Ind. Eng. Chem. Process Des. Dev. 1983, 22, 104-1 18
104
Process Slmulatbn of an SRC-I1 Plant Chandra
P. P. Singh‘ and Norman L. Carr
QuM Research 8 Development Company, PMsburgh. Pennsylvania 15230
This paper dsscribes the devebpment of simuiation of an SRC-I1 plant. The simulation considers all the significant units in the process flow sheet and also contains independent computational units to represent each combination 01 splkthg of process streams. Each unit model is developed independently and then u nW in the overen &mutation flow sheet, containing two recycle loops, to obtain the plant simulation. The simulation requires definition of only six independent variables to yield the specific Row rates and compositions of an the streams. The effect of changes in the values of the important variables on total liquld production rate (per unit volume of reactor) and liquid yieid (weight percent moisturdree coal) is evaluated. I t is shown that by making relathely small changes in tt.le operating condttkns, the liquid p~oductionrate can be increased by m e than 40% over those normally achieved. Difficulties in achieving complete conversion of liquefiable organic matter In coal are discussed.
Introduction Figure 1shows a simplified flow diagram of the SRC-I1 process considered in the present simulation. Coal and recycle slurry are mixed in a mix tank. The mixed slurry is heated to the feed temperature in a preheater and fed to the reactor. The preheater is not shown in Figure 1 since the rate-controlled reactions in the preheater are considered to take place in the reactor section, and the temperature of the feed to the reactor is not treated as a variable in this work. All the reactions are considered to take place in the reactor. Products from the reactor are separated into vapor and liquid streams in a high-pressure hot separator. Gases, water, most light distillates, and some middle and heavy distillates go in the vapor stream. All of the insoluble matter, SRC, and portions of heavy, middle, and light distillates are continued in the bottoms stream. The bottoms from the high-pressure hot separator are flashed in an atmospheric (50 psi) flash column, the bottoms of which are fed to the recycle splitter. The recycle splitter divides the bottoms from the atmospheric flash column into a recycle stream and feed to the vacuum flash column. Feed to the vacuum flash column also contains recycled light products from the same column. Light products from the hot separator and atmospheric flash column constitute the feed to a fractionator which yields light and middle distillates in the overhead stream, and liquids boiling above 270 OC in the bottoms. Vacuum column overhead product stream is added to the fractionator bottoms stream to make the “process solvent” or heavy oil stream, a part of which is recycled. Recycle oil and recycled bottoms from the recycle splitter constitute the recycle slurry. This simulation uses the SRC-I1 classification of the liquefied coal products which are numbered for the subsequent description of the unit models as: (1)inorganic mineral matter (ASH); (2) insoluble organic matter (IOM); (3) solvent-refiied coal (SRC); (4) heavy distillate (HD); (5)middle distillate (MD); (6) light distillate (LD); (7) water (H,O); (8)byproduct gases; (9) C1-CI gases, and (10) H2. The simplified process flow diagram (Figure 1) involves several interconnecting streams between various units and mixing and splitting of several process streams. Independent computationalunits are included in the simulation flow diagram (Figure 2) to evaluate the flow rates and compositions of streams resulting from mixing or splitting of various process streams. Also, all the units and streams are numbered as shown. The overall simulation program is organized in a modular fashion in which the unit numbers represent the se0 196-4305/83/ 1 122-0 104$0 1.50/0
quence in which the unit models are called by the main program for computation. Thus,the computations for any given set of process conditions start with the mix tank (unit 1) model by assuming the composition of the recycle stream. At the completion of one sequence of calculations (unit 141, a computed composition of the recycle stream (14-3) is obtained. This is matched with the assumed recycle stream composition to obtain correct values of flows and compositions throughout the plant. Within the overall recycle loop (from 1to 14), there is a smaller recycle loop involving units 6, 7, and 8 (Figure 2). The simulation determines the flows and compositions for this recycle loop, by an iterative process, before calling the model for the next unit, i.e., unit 9. After validation against extensive pilot plant data, the simulation is used to study the effeds of changes in the values of the independent variables on liquid yield (wt % mf coal) and liquid production rate (per unit reactor volume). Basis for Selecting Conditions for Study Several criteria such as productivity, i.e., amount of liquid (oil) produced per unit volume of the reactor per unit time, selectivity of a particular component, or consumption of hydrogen, operating cost, etc., may be utilized in selecting the best operating conditions for each unit. However, the capability of this analysis is limited. The kinetic model used to describe the rate of conversion of coal to different products describes the process by an instantaneous dissolution of coal followed by a reaction of SRC only, and the distribution of products is not affected by process conditions (Singh et al., 1982a). Therefore, selectivity of various components cannot be defined independently of each other. Once the amount of SRC reacted is known, the amount of each of the other components produced or consumed (hydrogen) is also known. Thus, selectivity in the present case is completely defined once the liquid yield (wt % mf coal) is determined. Therefore, we attempt to identify the operating conditions which influence the amount of liquid produced as well as liquid yield per unit weight of coal. It is expected that this study would provide a qualitative understanding of the relationships between operating conditions, liquid productivity, and liquid yield and a basis for quantitative evaluation of the optimum operating conditions. However, optimization per se was not attempted in this study. Models for Various Units Mix Tank Plus Preheater. The model for this combined unit evaluates the composition of the feed to the reador from known composition of recycle slurry, feed rate, 0 1982 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 105 OASES & WATER LIGHT 6 MIDDLE DISTILLATE
h
HYDROQEN
P
I
w
I
I
RECYCLE SLURRY
I
PRESS
IREACTOR
FRACTIONATOR
8EP.
4 D
DISTILLATES (1
-a)D
40
HEAVV PRODUCT0
Figure 1. Schematic flow diagram of SRC-I1 plant.
and mass fraction of coal in the feed. The calculated composition of the outlet slurry from the preheater includes the effect of initial coal dissolution and is expressed as Ggfi = G g f d i m d + RgiR
(i = 1, ..., 10)
(1)
where G is the total mass feed rate in kg/h (The total feed rate also includes input hydrogen added at the rate of 4 kg/lOO kg of feed slurry.), gfiis the mass fraction of component i in feed to the reactor, g f d is the mass fraction of coal in feed slurry, R is the recycle slurry rate in kg/h, g? is the mass fraction of component i in the recycle slurry, and fjd is the fraction of component i in the initial coal dissolution product. Values of fid for Powhatan No. 5 coal (10.1 wt % ASH) are listed in Table I. For other similar coals with a different ASH content, values of the distribution coefficients for initial dissolution are calculated from the equation (fi""')c = [(l - fimd')C/(1
- fi"')Powl(fi~~)Pow(2)
However, unless specified otherwise, ASH content of the coal is considered to be 10.1 wt % mf coal. Reactor. The reactor is assumed to operate isothermally under f d y backmiied conditions. The only reaction in the reactor is that of SRC reacting to lighter components. The rate of the SRC reaction is given by (Singh et al., 1982a) -rSRC
=
1.567 X 10' eXp(-79 16o/R,'I?p~:~~g~s~, kg/(Lh) (3) where R is the universal gas constant in kJ/(kmol K), pH is partid pressure of hydrogen in the reactor in MPa, ?1 is absolute temperature in K , and g, is mass fraction of ash in the reactor slurry. It is important to note that the
Table I. Distribution of Components in the Products of Coal Dissolution and SRC Conversion fraction from fraction from diss of coal, SRC conv, components wt % of mf coal wt % of SRC ASH IQM SRC HD MD LD water byproduct gases C,-C, gases hydrogen total
10.10 4.78 53.17 6.57 4.15 2.65 5.52
-0.09 100.0
23.21 35.79 16.87 3.85 1.25 35.45 -16.42 100.0
rate of SRC reaction is zero order; therefore, a possible variation in the concentration of SRC along the reactor length due to partial backmixing does not affect the rate of reaction. Also, due to high molecular weight of the product molecules, the change in mole fraction of hydrogen in the gas is only 20 to 30%. Since the rate depends on the 0.28th power of pH2(eq 3) and hydrogen gas is fed at more than one location along the length of the reactor, the fully backmixed assumption with an average partial pressure of hydrogen is considered reasonable. Variations of temperature inside the pilot plant reactors are observed to be small. This is also expected to be the case for larger, demonstration-plant or commercial-sizereactors. This is due to high levels of backmixing, high heat capacity of gas (Singh et al., 1982b), and introduction of quench gas at several locations along the length of the reactor to maintain a uniform temperature. Therefore, the reactor can be considered to be isothermal. However, gas holdups in the reactors of different lengths are expected to be significantly different since gas velocities are proportional to length for
106
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983
t
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983
a given space time, and gas holdup increases with increase in gas velocity. The rate of reaction is described in terms of nominal slurry residence time based on feed conditions (7 = VRps/G). Consideration of vapor-liquid equilibria and gas holdup in the reador could yield the actual slurry residence time. A rigorous analytical treatment of the process requires the use of actual slurry residence time. Although slurry properties, vapor-liquid equilibrium constants, etc., at the reactor conditions are not known accurately, a reasonable estimate of the actual slurry residence time can still be obtained. However, the kinetic model used in this work was based on the nominal slurry residence time (Singh et al., 1982a) at low gas holdups. Since a fairly large data base was used to develop the kinetic model, it is expected that the effects of changes in the amount of vaporization, due to changes in operating conditions, are absorbed in the model through the temperature and pressure dependence terms of the rate of SRC reaction (eq 3). Therefore, to use vapor liquid equilibrium relationships in order to calculate the actual slurry residence time, the kinetic model must be modified. This is beyond the scope of this work. The fraction of feed to the reactor converted to vapor phase is not expected to vary much. This is due to the fact that the feed contains 25 to 35 wt % coal, and the recycle slurry (65 to 75 wt% of feed) contains a very small fraction (about 5 wt %) of the distillates, expected to be transferred to vapor phase. The concentration of these low-boiling fractions also does not vary significantly with the operating conditions. However, for the sake of verification, variation in the effect of evaporation with changes in operating conditions has been studied by using the overall plant simulation and found to be very small. This is discussed in the Results and Discussion section of the text. This work includes the effect of gas holdup to express the rate of reaction in terms of reactor volume occupied by slurry. For this purpose, gas holdup (egs) in the experimental setup used to generate the data base for kinetic modeling work is estimated from the following correlation of Calderbank (1958) = E( a0
+ 0.015~
(4)
where c is holdup of gas in a stirred vessel, ugis superficial gas velocity in cm/s, ut is bubble rise velocity which is assumed to have a constant value of 26.5 cm/s, and gasliquid interfacial area, uo, is given by (5)
107
the liquid in cP. Gas holdups in the bubble column SRC-I1 reactors were obtained from the following empirical correlation of Kara et al. (1982) which has been developed for a three-phase system ug
tg =
Ai
+ Azug + A ~ U + L Aq[ts/(ts + EL)]
(8)
where tgis the gas holdup, uf is the superficial gas velocity, t, and tL are holdups of solid and liquid phases, and A's are parameters whose values depend on the size of the solid particles. For 3 - ~ m particles these are: A, = 13.23, A2 = 3.27, A3 = 2.39, and A4 = 87.83. In SRC-I1 reactors, eq 8 may be considered appropriate. The values of gas holdup obtained by use of eq 8 are comparable with those of other workers (Akita and Yoshida, 1973; Deckwer, 1980). With correction for gas holdup variations, the mass fraction of component i in the reactor or the outlet stream is obtained from the following component mass balance equation.
or
where gi is the mass fraction of component i in the reactor, is the fraction of component i in the SRC reaction product, VR is the volume of the reactor in L, and 7 is slurry residence time in h. Values of are listed in Table I. Separator Units. In the simplified flow diagram considered in the present simulation (Figure l),there are four separation units, namely the high-pressure hot separator, the atmospheric flash column, the vacuum flash column, and the fractionator. Except the fractionator, each of the separator units is essentially a flash column. Differences in their operating temperatures and pressures are determined by the quality of the feed they handle. However, irrespective of these differences between the three flash columns, the same algorithm (flash calculation procedure) is used by the models for these units. The description of the procedure of flash calculation used by these models follows. When the flash temperature and pressure are specified such that a two-phase mixture exists, the following overall and component material balances apply at steady state. F = VF LF (11)
fisRc
ftRC
+
Fxj =
VFyFi
+ LFxF~
(12)
"
which holds at moderate stirrer speeds such that
L
xxi = 1 ill
when this term exceeds 20000, the interfacial area a is given by = 1.95 x 10-5
(7) In eq 4 through 7, P, is the power dissipated per unit volume of liquid in hp/ft3, ut is the rise velocity of bubbles in cm/s, a is the surface tension in dyn/cm, dh is impeller diameter in cm, pL is the density of the liquid in g/cm3, N is the impeller speed in rpm, and pL is the viscosity of
where F is the rate of ASH and IOM-free feed to the column in kmol/h, xi is the mole fraction of component i in the feed, VF and LF are flow rates of vapor and liquid phases after flash in kmol/h, yFi and x F ~are mole fractions of component i in vapor and liquid, respectively, and C is the total number of components in the separator. Since the flash is considered to be an equilibrium vaporization
108 Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983
Table 11. Pseudocomponent Characterization no. main stream no. split stream 1. hydrogen 2. byproduct gases (CO, CO,, H,S, etc.) 3. water 4. C,-C, gases
5. LD ((2,-193 "C) 6. MD (193-288 "C) 7. HD (288-482 "C)
8. SRC (482+ "C)
1. hydrogen 2. byproduct gases (CO, CO,, H,S, etc.) 3. water 4. c, gas 5. C,-C, gases 6. C5-149OC 7. 149-193°C 8. 193-260°C 9. 260-288°C 10. 288-316 "C 11. 316-343 "C 12. 343-371 "C 13. 371-399 "C 14. 399-427 "C 15. 427-455 "C 16. 455-482 "C 17. 482-510 "C 18. 510+ "C
Table III. Split Stream Fractions main stream split stream G-C, LD MD HD
SRC
1-
i)])
(18)
Solution of the above equation, obtained by using the regula falsi iterative method, yields the value of LFIF which is used in eq 16 and 17 to obtain the composition of vapor and liquid phases. In the above discussed flash calculations, narrower boiling ranges of the coal liquefaction products are employed to get a better description of the separation. This clasaification is listed in Table 11. The distribution of the feed streams into narrow boiling range pseudocomponents is obtained from the true boiling point curve obtained by a blending computer program. The true boiling curve obtained in terms of volume fraction is used to obtain the composition of the distillate in terms of mass fraction (Table 111) r
normal bP, "C
177 23 2 26 8 302 329 357 385 413 44 1 468 496 552
(20)
where mi and mi are average molecular weights of components i and j , respectively. The values of average molecular weights for various pseudocomponents have been taken from the SRC-I1 design package (Process Design Specifications, June 1980) and are listed in Table IV. To obtain the compositions of flashed vapor and liquid streams, in terms of mass fraction, the following conversion is used C
gi = (mjxi)/ Cmjx, j= 1
(16)
Using eq 14 and 17, we have
E { % / [ 1-(;
= b / m i )/ Cgj/mi j=1
where Ki is the equilibrium vaporization constant for component i. Combining eq 11, 12, and 13, we have
1 = i=l
sp gr at 15°C 0.03 0.80 1.00 0.30 0.5100 0.7500 0.8914 0.9689 0.9881 1.0122 1.0328 1.0679 1.0944 1.1168 1.1375 1.1570 1.1752 1.2177
c xi
process, mole fractions of component i in the vapor and liquid phases are correlated by YFi = KiXFi
Various
where gi, ui, and pi are the maas fraction, volume fraction, and specific gravity of component i, respectively. For the purpose of flash calculations, the composition of the feed stream to any separator has to be described in mole fraction. This is effected by the following equation
mass fraction 0.25 ((2,-C,) 0.75 (C,-C,) 0.467 (LD) 0.533 (LD) 0.631 (MD) 0.369 (MD) 0.2 (HD) 0.226 (HD) 0.196 (HD) 0.146 (HD) 0.103 (HD) 0.064 (HD) 0.065 (HD) 0.143 (SRC) 0.857 (SRC)
Cl C2-G C,-149 "C 149-193 "C 193-260 "C 260-288 "C 288-316 "C 316-343 "C 343-371 "C 371-399 "C 399-427 "C 427-455 "C 455-482 "C 482-510 "C 510+ "C
Table IV. Physical Properties by Pseudocomponents av component mol wt hydrogen 2 byproduct 35 water 18 16 c, 44 C2-G C,-149 "C 100 149-193 "C 130 193-260 "C 158 260-288 "C 173 288-316 "C 185 316-343 "C 203 343-371 "C 220 371-399 "C 24 1 399-427 "C 260 427-455 "C 282 300 455-482 "C 482-510 "C 31 8 510+ "C 500
(21)
In all the flash calculations, IOM and ASH are solids and, hence, are retained in the slurry (liquid) phase. Therefore, C in eq 13 through 15 and 18 through 21 represents the number of volatile components, and mass, mole, and volume fractions are on the ASH and IOM-free basis. The flow rate and composition of the vapor phase remain unaffected by this change in basis. However, the actual flow rate of the liquid phase after separation is obtained by adding the flow rate of ASH and IOM to the flash calculated values. The actual mass composition of the slurry is determined from the equation
,.
L
g, = hxFimi/[CLFxFimi + WASH + EIIOM)~ (22) r=l
1. High-pressure Hot Separator. Experimentally determined values of the equilibrium vaporization constant, K , under the conditions existing in this separator for the SRC-I1 process are not available. Therefore, the values of K for the H-coal process, determined experimentally and correlated theoretically by the multiparameter corresponding states concept (Starling et al., 19801, are used. These values are obtained for conditions similar to those obtained in this separator and are available for all the 18 volatile pseudocomponents listed in Table 11. The adjusted values of K for all the components under the operating conditions prevailing in this separator are listed in Table V. For any specified set of conditions and known values of K , the model for this unit does flash calculations
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 100 Table V.
K Values for the Three Flash Separators
hot component separator H2
byproduct water
c,
C*-C, C,-149 "C 149-193 "C 193-260 "C 260-288 "C 288-316 "C 316-343 "C 343-371 "C 371-399 "C 399-427 "C 427-455"C 455-482"C 482-510"C 510+ "C
7.88 2.00 2.40 4.00 2.20 0.5225 0.3637 0.23 0.151 0.109 0.078 0.0525 0.035 0.021 0.0125 0.0075 0.0034 0.0005
intermediate vacuum separator flash column
311 51.10 26.54 159.0 55.4 8.54 2.54 0.94 228.37 0.43 141.06 0.25 84.17 0.144 50.31 0.081 29.03 0.0451 16.60 0.0243 9.18 0.0125 4.91 0.0061 2.50 0.0028 1.19 0.000132 0.1068 ~~
conditions pressure temperature
13.8 (MPa) 345 ( e a ) 427 "C 390 "C
172.5 (Pa) 335 "C
to yield flows and compositions of vapor and liquid phases. 2. Atmospheric Flash Column. Correlations for predicting the values of K for the 427-510 "C boiling range fractions have been recently reported (Gopal et al., 1981). These correlations have been developed for vacuum flash column operating conditions. These values are corrected for pressure, and the K values for lower and higher boiling range components are obtained by a linear extrapolation of the form log ( K ) = d + 6 T b (23) where T b is the average boiling temperature for any pseudocomponent. K values for this column under a typical set of conditions are listed in Table V. The values of K being known, the model for this separator evaluates the compositions and flow rates of the product streams. 3. Vacuum Flash Column. The K values for components in the temperature range of 427-510 "C are available, and for other pseudocomponents the value is correlated by eq 23. These values are listed in Table V. In the present study, the amount of overhead recycle to feed has been assumed to be 0.6 times the net feed to the column. 4. Fractionator. This unit is considered to split its feed into two streams. The overhead stream contains all liquids boiling below 270 "C, and the bottom stream contains those boiling above 270 "C. This corresponds to 75 w t 7% of the middle distillate and the lighter components going to vapor phase (Table 111). Stream Mixers. Units 6, 10, 12, and 14 (circled in Figure 2), namely, fractionator feed, vacuum feed, liquid streams mixer, and recycle slurry mixer, respectively, are stream mixers used as independent units for computational purposes only. The inlet streams to all these units are numbered 1and 2, whereas number 3 represents the outlet streams. Flow and composition of the mixing product stream are determined by the unit model from the inlet streams flows and compositions as follows G (m,3) = G (m,l) + G (m,2) (24) and gi (m,3) =
G (m,l) gi (m,l) + G (m92) gi (m92) G( ~ 3 ) (i = 1, 2, ..., 10) (25)
where m and 1,2, or 3 in parentheses represent unit and stream numbers, respectively. Stream Splitters. There are three stream splitters involving atmospheric flash bottoms, recycle or heavy distillate, and vacuum flash overhead. Each of these splitting of streams is for separating a liquid or slurry stream, after removal of gases and lighter products, into a recycle and a product stream. Recycle bottoms splitter is a real unit in a SRC-I1plant, whereas liquid recycle and vacuum overhead splitters are computational units used for simulation purposes (Figure 2). Compositions of the split streams are the same as those of the feed stream, and flow rates in the split streams are defined by an independent variable. Therefore, the stream splitter models essentially assign the composition of the feed stream to the resulting streams and divide the flow rates in the preassigned ratios as follows
G (m,2) = a (or 0)G (m,l)
(26)
G (m,3) = (1- a (or 0))G (m,l) (27) (i = 1, 2,..., 10) (28) gi (m,3) = gi (m,2) = gi (m,l) where a or is the ratio in which feed stream is split. It is important to mention here that the recycle bottoms splitter (unit 5 in Figure 2) consists of two compartments, the outlets from which are the two resulting streams. Recycle of the overhead from vacuum flash separator is fed to the compartment which feeds the vacuum flash unit. Mathematically, the real arrangement is the same as the one shown in Figure 2. Overall Plant Simulation In addition to all the unit models described in the preceding section, the overall plant simulation requires the definition (numerical values) of the exact number of variables required to satisfy the degrees of freedom for the system. To define the degrees of freedom for the plant, overall and component material balances around the plant are considered. For the sake of convenience, the following discussion of the overall simulation is confined to the simplified flow diagram in Figure 1. The separator units are numbered as: (1) high-pressure hot separator, (2) unified separator, representing separation of gases and water from the overhead stream high-pressure hot separator before it goes to fractionator, (3) atmospheric flash column, (4) fractionator, and (5) vacuum flash column. Overall mass balance (Figure 1) L, + L4 + H5 + (1- a)D (29) Gg+,,, Component mass balance (Figure 1)
Lzgb
+ L&+
+ (1- a)DgiD
HgF
(i = 1, ..., 10) (30)
10
cgfi = 1
i=7
10
Cg? = 1 i=4
(i = 1, 2, 3)
gk' = 0
(33) (34)
R
kgi" = 1 i=4
giD = 0
(i = 1, 2, 3, 7, ..., 10)
(35) (36)
110
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 4
CgrH5 = 1
r=l
(37)
gLH5= 0 (i = 5, ..., 10) (38) where L2 and L4 are flow rates of vapors from separators 2 and 4, respectively in kg/h, H5is the vacuum flash bottoms flow rate in kg/h, D is the total distillatea in kg/h, gls are mass fractions of component i in the superscripted streams, and V Ris the volume of the reactor in L. From eq 3 and 29 through 38, it can be noted that there are 52 variables (40 gi)%G, g f c d VR, up PH2, T , gASH, L2, L4, H5, D , and CY),whereas there are only 37 independent equations. If all the variables were independent, the degrees of freedom (= number of independent variables - number of independent equations) would be 15. Therefore, definition of (assignment of numerical values to) any set of 15 independent variables would uniquely define the flows and compositions of all the streams in the plant (Figure 1). It is important to note that degrees of freedom of 15 also means that there are only 15 independent variables in the system. However, the total number of variables can be much larger. For example, in a set of variables G, VR,and 7,only two are independent; the value of the third is defined by the following equation. 7 = (v~p,)/G (39) In this study, variations in the operating conditions of the four separators, namely, high-pressure hot separator, atmospheric flash column, vacuum flash column, and fractionator, are not considered. Therefore, eight variables (pressures and temperatures in each of the separators) have only one value, and, in general, these eight degrees of freedom are not utilized for variable study purposes. Also, production per unit volume rather than total production for a given reactor size is the basis for discussion. Therefore, slurry residence time (T),reflecting the ratio of reactor volume and volumetric flow rate, is enough to define the system. Thus, in the following sections the degrees of freedom are considered to be six (=15 - 9) only. Unless specified, the magnitude of superficial gas velocity is evaluated by considering hydrogen feed rate to be 4 w t 9% of the slurry feed rate and purity of the feed gas is considered to be 90 mol %. Procedure of Calculation Seven variables, G, V,, g,, g f d ,p H , T, and ug,are fixed to define the system. The recycle how rate is obtained from the following overall mass balance around the mix tank. R = G(1 - g f m d (40) Solids mass fraction in the recycle stream klR+ g2R)is obtained from the equation
RkiR+ g2R) = G k s - gfcod
(41) Mass fractions of ASH and IOM in the recycle stream are determined by using the fact that the ratio of the two in any stream is the same as their instantaneousdissolution product ratio from coal, i.e. (42) Mass fractions of the other four components (SRC, HD, MD, and LD) in coal are assumed as follows g5R =
g,R = 0
(43)
(44) g3R = g,R = (1- glR - g 3 / 2 With this assumed set of values for the recycle slurry
composition, the main program starts the calculations by calling unit models according to their sequence numbers in Figure 2. For evaluating flow to the vacuum feed unit 6, the recycle ratio, j3,is obtained from the following mass balance equation (Figure 1)
P = 1 - GgfcoafiCoal/&iB
(45)
Recycle ratio of liquids for the liquid recycle splitter is obtained from a = ( R - @)/D (46) Evaluated composition of the recycle stream is compared with the values from the previous iteration. If the difference for either of the components is more than 0.1% , i.e.
then the calculated composition of the recycle is used to repeat the iterative process. Results and Discussion A. Effect of Evaporation in the Reactor. Values of the vaporization equilibrium constant ( K ) ,under the reactor conditions, are not available. These have to be obtained by extrapolation of the corresponding values for high-pressure hot separator conditions. The K values for the high-pressure hot separator, listed in Table V, are also not general enough to claim their applicability to all coals or their meaningful extrapolation to different sets of pressure and temperature conditions. However, the magnitude of K values decreases by five orders of magnitude as the boiling range of liquids increases from c5-149 to 510+ "C. Therefore, it is expected that a small error in the magnitude of K values used in this work (Table V) should not have any significant influence on the predicted extent of vapor-liquid separation. However, the K values are very sensitive to temperature, and the effect of temperature needs to be considered for use of the hot separator K values for vapor-liquid separation in the reactor. For this purpose, it is assumed that change in the K values of a component is proportional to change in its vapor pressure. Measured values of vapor pressures of the pseudocomponents at various temperatures have been correlated by the following general expression. b2i
In (VP,) = bIi - (48) T where VPiis vapor pressure of component i at temperature T (K)and bli and bzi are parameters whose values are estimated from experimental data. Based on the above correlation, the K value for a component i at any temperature, T,is expressed as -
I
\
-
where (KJ0is the known K value for component i at temperature To. Values of bzi for all the pseudocomponents are listed in Table VI. As mentioned earlier, the reliability of the K values listed in Table V is rather limited. In order to see the effect of errors in K values, the latter is varied over a wide range ( K / K o = 0.6 to 1.5) to generate Figure 3. Five selected cases representing large variations in process conditions and hence the yields are considered (Table W). Figure 3 shows the effect of vaporization in terms of the weight percent of feed slurry remaining in liquid (or slurry)
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983
Table VI. Values of the Parameter ( b z i )for Vapor Pressure Correlations component i boiling range C,-149 "C 149-193 "C 193-260 "C 260-288 "C 288-316 "C 316-343 "C 343-371 "C 371-399 "C 399-427 "C 421-455 "C 455-482 "C 482-510 "C 510+ "C
parameter ba ( e s 49) 4300 4900 5500 6100 6500 7100 8070 8200 8440 8700 9000 9400 10000
ARE L I S T E D I N TABLE- VI1
Table VII. Operating Conditions and Yields for Figure 3 case no. gf coal
gs pH T, ' C'MPa 7,
h
ug, cm/s
I I1 I11 IV V 0.35 0.30 0.25 0.30 0.35 0.50 0.45 0.45 0.50 0.50 11.04 11.04 11.04 11.04 11.04 455 445 445 465 465 1.11 1.11 1.0 1.11 1.0 1.0 1.0 8.2 1.0 1.0
yields, wt % mf coal liquid
41.7
29.1
30.5
45.8
59.6
SRC hydrogen
25.0 -4.1
41.4 -2.0
39.5 -2.3
19.6 -5.5
1.4 -8.2
(HD+ MD + LD)
phase in the reactor. For the first four sets of conditions (cases I through IV), the maximum difference between extents of vaporization are within 5 wt % at any particular value of K or KIK,. This corresponds to a maximum slurry residence time variation of about 6% or &3%. It is important to note from Table VI1 and Figure 3 that the extent of vaporization in the reactor is also a function of liquid yield. However, for very high liquid yield represented by case V, evaporation effect at high K values (KIK,, = 1.5) is significantly higher. This can result in about 30% larger slurry residence time for case V compared to other cases. However, case V also represents only a hypothetical situation, since the plant would become inoperable, due to an excessive buildup of solids, much before achieving almost complete conversion of liquifiable organic matter as in this case. A 3&46 wt % (mf coal) liquid yield range covered by conditions in cases I through IV may be considered to be large enough to provide a general assessment of the evaporation effects in the reactor. Based on cases I through IV, it can be concluded that evaporation in the reactor under different seta of operating conditions can lead to a difference of about 5 wt % in the amount of evaporation in the reactor. If the effect of variations in composition of the slurry is assumed to change the K values by &lo%, evaporation in the reactor is expected to affect the actual slurry residence time by less than *4%. In view of limited accuracy of the available experimental data, such a small variation in the actual slurry residence time is considered to be insignificant and the reaction rate expression based on nominal slurry residence time ( 7 ) (eq 3) is used in this analysis. B. Validation of the Simulation. The kinetic model and reaction scheme were developed from specially designed kinetic experiments where the variable space was quite large. Now we can use pilot plant data for inde-
111
112
Id.Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 50
I
I
I
/
/,
' /
+ 1.5 / WT%
e;,
c~~= -
/
'
/-
~
i I
10
0 1 2 3 4 5 6 HYDROGEN CONSUMPTION (EXPERIMENTAL). WT% mf COAL
I
20 30 40 L I Q U I D YIELD ( E X P E R I M E N T A L ) , W T % m l COAL
50
Figure 6. Parity plot for hydrogen consumption.
Figure 4. Parity plot for liquid yield. 60
I
I
I
I
I
I
/
I/ 0 .
I
20
19.1% 16.7%
30 40 50 SRC YIELD (EXPERIMENTAL). WT% mf COAL
I 1 00
Figure 5. Parity plot for SRC yield.
parity plot for SRC yield does not provide an equally good validation for the simulation (Figure 5). Overall average and absolute errors are 18 and 21%, respectively. Although almost equal values for absolute and average errors do indicate that scatter in the data points around an average is small, the calculated SRC yields are on the average about 18% (-5 wt % ) higher than the measured values. However, it is very important to note that the data base contains many runs with Powhatan No. 5 coal (Table WI) which had been used to generate the kinetic model data base (Singh et al., 1982a). Although a detailed analysis to account for individual coals is not presented here, it was noted that there is no significant difference in the SRC yield behavior for the different coals. There are several reasons which may account for the average difference. These are as follows. 1. Magnification of Measurement Error Due to Evaluation of Yields by Difference. On the average, the feed slurry contains 30 wt % coal, whereas measurements are made for the total slurry. Therefore, an experimental error of f l % is magnified to i3.3% for the yields based on mf coal. 2. Transfer of Measurement Errors to Nonliquid Products. Measurement of the total liquid yield is gen-
erally more reliable than that of the other components; therefore, any loss or gain of material in the material balance is normally assigned to nonliquid products. Since the output is more frequently less than the input, heavy as well as gaseous products may be estimated to have lower values. 3. Basis for ASH Balance. ASH balance is used to obtain a better estimate of the yields. For this purpose it is generally assumed that the only loss from the feed inorganic mineral matter (ASH) is that of about half of the pyritic sulfur. However, in the kinetic modeling work it was noted that the measured ASH content of feed Powhatan No. 5 coal was 12.1 wt % (mf coal), whereas the corresponding measured value after reaction was only 10.1 wt % (Singh et al., 1982a). Since the products are classified into gases, liquid, and vacuum tower bottoms before their final classification into individual components, the assumption of no change in chemical composition of ASH can lead to substantial (>2 wt %) measurement error. 4. Conversion of SRC to IOM in Vacuum Flash Column. In the vacuum flash column, a small fraction of SRC is known to be converted to IOM and may be gaseous products. Since no vacuum flash separation was involved in the kinetic experiments (Singh et al., 1982a), the pilot plant data are expected to show a somewhat lower yield of SRC. From the above discussed possibilities, it seems more likely that the large average error and bias in the SRC yield parity plot may be due to the estimation of SRC experimental yield or due to the difference in separation schemes of the A-1 and pilot plant units. In the latter case, a correction needs to be applied to the plant simulation. The average measured SRC yield is about 15 wt % lower than the calculated values. However, due to lack of sufficient basis for such a modification as well as the basic interest in liquid yield, no modifications to the kinetic model are made at this stage. Absolute error for hydrogen consumption is the same as that for SRC yield, and average error is slightly smaller than the corresponding value for SRC (16 vs. 18% for liquid yield). This shows that scatter in hydrogen consumption data is relatively large (Figure 6). The almost equal average error for liquid yield and hydrogen consumption indicates that the transfer of measurement errors to nonliquid products is a more likely source of error. In view of the relatively large data base (99 measured sets of values), the simulation is considered to be validated, at least for the purpose of studying the effects of various
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 I I
I
'
I
I
I
'
I
'
'
113
I 1 +200
37t
go= 0.45
-
9fCoaI=0.30 1.108h T = 455.c P =13.8MPa u / = o . 1 ug
7
9s-045 O f c o a i = 0 30 T = 1 108h T 4 5 5 'c
-a
-
u
E
f
5
38
-
ug-8
15 cmla
2 'y
P D 35
r-
34 10
11
12
13
14
15
Preaaure (P). MPa
Figure 9. Effect of reactor pressure on liquid yield and liquid production rate.
I
351-
0
5
10
40L
15
Superficial G a s Veloclly (ug), c m l a
Figure 7. Effect of superficialgas velocity on liquid yield and liquid production rate, due to change in gas holdup. 40
1'
38b!
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I
n: I
I
Q f c o e I - 0 30
7 -1.108
-
T=455'c P 13.8MPa Ug-8 15cmla
'413.0
12.68
38
1
,
5
08 Ofcoal P-13.8 - 0 . 4-0.30 5U p s
f-l.lO8h
'P
12.35
-1:
N-
12.02
x
J
l
/
4
8.12
U
4
111.27
.!{
P 10.95
20 30
35
Solid8 445
450
465
480
485
40
In Feed Slurry
45
50
(98). Wt96
Figure 10. Effect of solids concentration in feed on liquid yield and liquid production rate.
Temperature (T).'c
Figure 8. Effect of reactor temperature on liquid yield and liquid production rate.
variables. The magnitude of the variables held constant in studying the effect of change in the value of one variable at a time is the same as those proposed for the SRC-I1 demonstration plant or generally found from the P-99or Ft. Lewis pilot plants. C. Effect of Superficial Gas Velocity, ur This effect is shown in Figure 7. Liquid yield decreases with an increase in superficial gas velocity due to an increase in gas holdup and a decrease in "actual" slurry residence time. (It is important to note that the "actual" term has been used only to differentiate the nominal slurry residence time, based on feed conditions, from that evaluated by considering the effect of gas holdup.) The rate of decrease in liquid yield decreases at higher gas velocities since the rate of increase in gas holdup or decrease in slurry residence time is lower a t higher gas velocities (eq 8). It is important to note that decrease in liquid yield with increase in superficial gas velocity is not proportional to the decrease in slurry residence time, since about 20 wt % of liquid is produced by instantaneous dissolution of coal and
is not subject to changes in the rate of SRC reaction. D. Effects of Pressure, Temperature, and Solids in Feed Slurry. Liquid yield as well as liquid production rate [kg/(L h)] increases with increase in value of either of the three variables, namely, pressure, temperature, and solids concentration in feed (Figures 8 through 10). In each of these cases, the increase in liquid production rate is due to higher rate of SRC reaction, and the liquid production rate [kg/(L h)] is proportional to liquid yield (wt % mf coal). The only other important point to note from these figures is the magnitudes of increase in the three wes. An increase in temperature from 445 to 465 OC,Le., 20 "C increase, results in more than 6 wt % increase in liquid yield (33.6to 39.6 w t ?& mf coal), whereas increase in pressure from 10 to 15 MPa results in a liquid yield increase of only about 2.4 wt %. A smaller effect of increase in pressure is expected from 0.28th power dependence of the rate of SRC reaction on hydrogen partial pressure (eq 3). The increase in liquid yield is proportional to increase in solids concentration in the feed slurry. For the conditions listed in Figure 10, the proportionality constant is 0.85 wt ?& liquid per wt % of solids in the feed slurry.
114
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983
Table VIII. Data Base for Validation of the Simulation f y l
no.
unit=
coalb
x 100
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60 61 62 63 64 65 66 67 68 69 70 71 72 73 74 75 76
P99 P9 9 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P99 P9 9 P99 P99 P99 P99 P9 9 P99 P99 P99 P99 P99 P99 P9 9 P99 P99 P99 P99 P99 P9 9 P99 P99 P99 P99 FTL FTL FTL FTL FTL FTL FTL FT L FTL FTL FTL FTL FTL MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER
POW POW POW POW POW POW POW POW POW POW POW POW POW POW POW POW POW POW POW POW KEN KEN KEN KEN KEN KEN KEN KEN IRE IRE IRE IRE IRE IRE IRE IRE IRE
12.2 12.2 12.2 12.5 12.5 9.1 9.1 9.1 9.1 11.5 11.5 11.5 9.5 11.5 11.5 11.5 11.5 11.5 10.7 10.7 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 11.3 11.3 11.3 11.3 11.3 12.8 15.6 16.3 16.3 13.1 16.5 16.5 16.5 16.5 11.5 11.5 11.5 11.5 11.5 11.5 11.5 11.5 9.5 9.5 9.5 9.5 9.5 11.5 11.5 11.5 11.5 11.5 11.5 11.5 11.5 11.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5
BLK RR RR RR RR POW POW POW POW POW POW POW POW KEN KEN KEN KEN KEN POW POW POW POW POW POW POW POW POW KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN
gf coal
x 100 29.4 29.7 29.5 30.5 30.3 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 10.1 15.2 15.1 22.2 22.2 30.0 30.0 33.2 29.6 32.8 32.7 29.8 29.6 29.6 29.7 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 29.4 29.5 24.5 24.4 29.2 27.1 27.9 30.9 30.2 30.6 31.3 29.6 26.1 24.2 24.6 29.9 33.5 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 35.0 40.0 45.0 45.1 40.1 35.0 35.0 40.0 45.0 35.0 30.0 45.0
T,K 455 455 455 455 455 4 54 4 54 458 457 458 4 53 4 58 455 44 0 444 44 3 444 44 7 451 4 52 460 460 46 5 460 460 460 460 460 460 458 460 460 452 4 58 460 457 457 455 455 455 455 455 4 56 4 56 456 46 0 4 56 4 56 455 455 461 4 59 459 461 4 59 457 455 455 455 455 455 455 455 455 461 455 455 455 455 455 4 54 463 455 455 455 455
PH
MA
9.02 9.07 9.25 9.37 11.72 9.37 10.36 9.37 10.36 10.36 10.36 10.36 9.37 10.64 10.95 11.17 10.89 10.52 9.82 9.82 11.75 13.02 11.65 12.76 11.59 14.60 14.70 11.77 9.76 9.82 10.14 9.82 9.84 9.82 9.37 10.36 10.36 9.37 9.84 8.94 9.65 8.93 11.81 11.81 11.87 11.69 11.85 11.80 11.86 11.48 12.49 12.03 12.15 11.93 12.04 12.51 12.51 12.51 12.51 12.51 12.51 12.51 12.76 12.90 13.89 13.20 13.20 13.20 13.20 13.20 13.20 13.20 13.20 13.20 13.20 13.20
7,
h
1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.95 0.95 1.00 1.00 1.00 1.00 1.33 1.33 1.00 1.00
1.00 1.06 1.00 1.04 1.33 0.66 0.66 1.22 1.21 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.85 0.85 1.00 1.00 1.00 1.00 1.00 0.91 0.86 0.88 0.97 0.95 0.97 0.97 0.93 1.00 0.98 0.82 0.80 0.96 1.01 1.00 1.01 1.02 1.01 1.02 1.01 1.01 1.05 0.96 1.01 1.04 1.49 1.59 1.02 0.98 1.08 0.99 1.04 0.97 0.99
gs x
100
43.8 43.0 42.6 44.2 44.8 44.0 42.0 42.0 42.0 45.0 44.0 45.0 43.0 17.2 22.8 23.4 33.4 33.4 45.5 44.7 44.1 43.2 41.9 40.6 42.4 45.4 43.4 42.6 44.0 44.9 47.0 45.0 45.2 45.0 42.0 46.0 48.0 43.0 45.4 46.8 45.4 45.5 44.4 43.3 44.8 47.4 43.9 43.8 45.2 43.5 39.8 39.3 41.1 42.7 44.3 42.1 41.3 41.9 43.6 44.0 46.0 43.8 43.6 44.3 47.9 48.1 52.8 53.1 50.6 46.5 46.9 49.0 55.6 46.8 44.7 48.0
33.0 33.5 32.8 36.5 41.3 30.8 31.1 29.9 30.4 40.2 37.8 39.7 37.3 31.4 33.8 36.6 33.9 35.8 38.1 35.9 37.4 41.6 38.2 35.2 33.1 43.1 44.3 41.1 37.9 39.6 41.9 39.4 39.7 31.0 34.4 35.4 38.9 29.0 33.2 33.9 36.3 36.9 43.3 41.2 40.7 37.1 43.2 45.5 44.1 40.0 42.5 49.7 45.6 37.4 31.7 44.1 43.5 44.5 39.5 41.3 43.4 44.9 42.7 45.7 44.0 35.4 27.6 27.8 39.0 35.5 34.3 31.4 27.7 35.0 44.3 27.8
hSRC
hHz
28.0 28.7 31.1 23.2 19.2 29.0 31.7 28.9 32.6 22.4 26.3 22.8 25.7 33.9 31.1 27.1 30.7 28.9 23.1 24.9 27.6 27.4 25.9 27.5 32.2 27.5 18.7 19.2 24.8 22.4 24.2 24.2 23.6 24.0 22.1 22.4 20.3 28.7 23.9 24.9 21.3 20.6 30.7 27.7 25.6 27.7 30.4 26.2 28.2 31.6 25.7 22.9 23.3 28.8 37.9 28.3 30.9 27.7 30.8 29.9 26.0 28.2 29.7 28.8 23.5 37.8 46.0 32.2 29.6 37.8 33.8 36.9 44.2 36.7 27.3 45.9
-3.80 -3.80 -3.50 -4.20 -4.70 -3.30 -3.30 -3.50 -3.30 -4.00 -3.80 -4.00 -4.30 -4.00 -3.70 -4.30 -3.90 -3.70 -4.10 -4.00 -5.42 -5.15 -5.24 -5.52 -4.42 -5.10 -6.49 -5.81 -4.10 -4.20 -4.40 -3.90 -4.10 -3.70 -4.20 -3.90 -4.00 -3.50 -3.70 -3.50 -4.00 -4.00 -4.50 -4.60 -5.50 -5.20 -3.90 -4.50 -4.50 -4.40 -4.53 -4.10 -4.50 -4.30 -3.60 -5.64 -5.41 -5.30 -6.03 -5.22 -4.65 -5.98 -5.30 -5.17 -5.21 -3.26 -3.68 -5.71 -4.99 -2.94 -3.33 -3.61 -3.14 -3.14 -4.12 -2.95
u€!,
cm/s 1.00 1.00 0.94 0.91 0.91 0.90 0.90 0.90 0.90 0.89 0.89 0.89 0.88 0.91 0.91 0.90 0.92 0.91 0.89 0.89 1.61 1.09 1.19 0.95 1.49 1.49 1.09 1.18 0.91 0.91 0.91 0.90 0.88 0.86 0.98 0.97 0.97 1.00 0.90 1.09 1.10 1.12 1.86 2.12 2.22 2.06 2.28 2.91 2.77 2.25 3.36 3.41 4.66 5.74 4.97 0.48 0.48 0.67 0.67 0.48 0.67 0.47 0.46 0.46 0.67 1.29 1.26 0.88 0.83 1.29 1.31 1.31 1.27 1.19 1.78 1.66
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 115
77 78 79 80 81 82 83 84 85 86 87 88 89 90 91 92 93 94 95 96 97 98 99
MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER MER
40.0 30.0 45.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0 40.0 30.0 30.0 30.0 30.0 30.0 30.0 30.0
9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 9.5 11.9 12.2 12.3 10.3 14.2 15.9 14.1
KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN KEN I+ B I+ B I+ B I+ B I+ B I+ B I+ B
455 455 455 466 455 455 455 455 455 455 455 455 454 455 463 463 457 457 450 467 457 457 457
13.20 13.20 13.20 13.89 13.89 13.89 14.48 11.72 13.89 13.89 13.89 13.89 13.89 13.89 13.89 13.89 9.93 9.93 12.42 9.93 9.93 9.93 9.93
1.00 1.00 1.48 0.68 1.02 0.99 1.03 1.06 1.04 1.00 1.00 0.99 1.07 1.05 0.133 0.13 1.03 1.01 1.02 1.02 1.01 1.00 1.01
33.1 38.5 37.0 41.3 42.8 39.2 45.5 39.0 42.1 43.1 36.4 41.2 44.0 43.5 16.5 21.3 34.1 32.6 35.2 35.9 35.1 33.0 35.0
48.0 43.9 54.5 48.7 48.7 49.1 47.0 48.4 49.3 49.2 34.9 44.8 47.7 45.9 32.9 43.2 44.5 45.0 43.7 43.7 44.5 43.2 45.6
1.71 1.81 0.93 2.22 1.02 0.81 1.16 1.44 1.22 1.22 1.22 1.22 0.44 0.44 0.90 0.90 1.07 0.97 0.87 0.90 0.87 0.88 0.77
-3.55 -3.97 -4.52 -5.02 -5.72 -4.86 -6.21 -5.93 -5.71 -6.20 -4.79 -5.35 -4.77 -5.99 -2.43 -1.77 -4.90 -4.40 -4.60 - 5.20 -5.00 -4.80 -4.40
38.6 34.3 30.8 22.9 23.5 26.6 21.8 30.3 27.3 23.3 40.5 35.7 25.1 31.1 65.0 59.6 27.8 29.9 30.6 28.7 26.7 27.5 25.3
Units: P99, SRC-11 Process Development Unit at GR&DC, Harmarville; FTL, SRC-I1 Pilot Plant a t Ft. Lewis; MER, Coals: BLK, Blacksville coal; IRE, Experimental (1-L) Reactor of Pittsburgh & Midway's Research Facility in Merriam. Ireland Mine coal; I+B. Mixtures of Ireland and Blacksville coals; KEN, Kentucky coal; POW,Powhatan coal; RR, Robinson Run. 2.40
1.71 I
Slurry Feed Rate. k9NL.h) 1.33 1.09 0.92
]
]
I
-
I
]
]
I
8
0.8 120.0
\
0.45
gfcoal -0.30 T-465Oc P 13.8MPa ug -8.15cmls
-
-
1
E f
I \
%
?=
1.108h T= 456%
\
1 1 1
34 I
\
I
2620
I 25
11.6
30
40
35
0.7
0.0
1.1
1.3
1.5
Coal Concentration in Feed (Ofcoal).
Slurry Residence Time ( T ) . h
Figure 11. Effect of slurry residence time on liquid yield and liquid production rate.
E. Effect of Slurry Residence Time. As shown in Figure 11, liquid yield increases linearly with an increase in slurry residence time, whereas liquid production rate decreases with an increase in slurry residence time. A decrease in liquid product rate [kg/(Lh)] with an increase in liquid yield results from the fact that n'ith increase in slurry residence time, the feed rate of coal per unit volume of the reactor is reduced. The loss in the amount of liquid produced by instantaneous dissolution of coal, due to reduced coal feed rate, is not compensated by relatively higher liquid yield (wt % mf coal). For example, for the set of conditions listed in Figure 11, an increase in slurry residence time from 0.5 to 1.5 h results in an increase of liquid yield from 27.25 to 41.7 wt % mf coal, whereas the coal feed rate decreases by a factor of 3. The net result
f
1i :=
\
30
0.5
11.8
P = 13.8MPa up-8.15 cmlr "E-0.90cmls
WtS
Figure 12. Effect of coal concentration in feed on liquid yield and liquid production rate.
is a significant decrease in the liquid production rate per unit volume of the reactor. It is important to note that such a relationship between liquid yield and liquid production rate results from the basic nature of the process consisting of instantaneous dissolution of coal followed by rate-controlled reaction of SRC. This interesting relationship between yield and production rate is discussed further in the following paragraphs. F. Effect of Coal Concentration in Feed. For the set of conditions listed in Figure 12, an increase in coal concentration in feed from 30 to 35 wt % decreases the liquid production rate by about 1 % , whereas a decrease in coal concentration, from 30 to 25 w t % ,in feed increases the liquid production rate by about 1 wt %. Although relatively simple to predict theoretically, it may be difficult
116
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 r
T
---T
---
3 - - 7
T-
-7
7-
9s - 0 . 4 5 r = 1 108 T=455% P-13.8MPa ug - 8 . 1 Scmlr ut-o.gacm/s
1
‘I
8
!
1
I
I
E
?q 7 0 B
Region of interest
20
c
i
I
I
React1 I
Coal Concentration in Feed (91coal). W t %
to comprehend this effect if it is not noted that an increase or a decrease in the concentration of coal in feed slurry is always associated with a corresponding decrease or increase in the concentration of solids in the recycle slurry such that a constant level of solids in the feed is maintained. In Figure 12, for 25 and 35 wt % coal in feed the liquid yields are 44.1 and 30.9 wt % mf coal, respectively. Thus, the liquid yields for two feed coal concentrations differ by about 13 w t %. However, the differences in liquid production rate [kg/(L h)] differ by only 2 wt % since the feed rate of coal for the 35 wt % concentration of coal in feed is 40% more than that in the case of 25 wt % coal in feed. Thus, the loss in liquid yield is more than compensated by an increase in the feed rate of coal. It is very important to note that increases in the feed rate of coal are able to compensate for loss in yield due to the fact that most of the liquid (about 20 wt % mf coal) is produced by instantaneous dissolution of coal which is not affected by varying the process conditions. If all of the liquid products resulted from the slow reaction of coal, or dissolved coal, i.e., the SRC, liquid production rate for 35 wt % coal in feed would have been about 30% lower than that for 25 wt % coal in feed. It is noted from Figure 12 that liquid yield, w t % mf coal, increases monotonically with a decrease in concentration of coal in feed; however, the liquid production rate shows a maximum with respect to coal concentration in feed at the latter’s value of 23.5 wt %. If the coal concentration in feed is decreased below 23.5 w t %, the oil production rate decreases rather rapidly. Although the overall effect of coal concentration in feed over a wide range of ita values (20-40 wt %) is rather small (3 wt % ), it is important to note that the pilot plant operating conditions or design conditions for demonstration plant do not represent the optimum value of coal feed concentration. Handling of slurry containing high mass fraction of solids and high hydrogen requirement (wt % mf coal) are two important factors which may limit the operation of an SRC-I1 plant to a lower than maximum production level. Figure 13 shows the variations in solids content of the slurry and hydrogen consumption at conditions identical with those in Figure 12. From this figure, it is noted
1
I I I/ /I I 0 8
Figure 13. Solids concentration in vacuum column bottoms and hydrogen consumption rate as a function of coal concentration in feed.
I
07
08
0.9
10
Recycle Ratio ( b )
Figure 14. Solids concentration in vacuum column bottoms and reactor as a function of recycle ratio.
that for operation of the plant at maximum liquid production rate [kg/(L h)] the mass fraction of solids in the vacuum tower bottoms is about 46 wt %. It is close to the maximum solids concentration that the vacuum flash column can usually handle. Hydrogen consumption level of 5.7 wt % mf coal is also very high. This high hydrogen consumption requirement cannot be met by gasifying vacuum column bottoms, containing SRC and IOM only. This is due to smaller yield of SRC for high liquid yield. Thus, if the liquid production rate is to be maximum, the vacuum flash column may have some operational problems, and additional hydrogen will have to be generated by using a source other than vacuum column bottoms. However, if it is required that almost all liquifiable organic matter in coal be converted to liquid products, this process offers the possibilities of achieving it. G. Sensitivity of Solids Concentration in the System to Recycle Ratio. In all the above discussed sets of conditions, solids concentration in the feed is assumed to be maintained at a certain level. This is effected by fixing the recycle ratio, p, from the recycle splitter. Figure 14 shows variations in solids concentration in the reactor and vacuum column bottoms. These two represent the minimum and maximum solids concentration levels in the plant, respectively. Concentration of solids in the highpressure hot separator and atmospheric flash column lie in between these two extremes. Based on the assumption that the maximum solids concentration in the vacuum flash column is about 45 wt % ,it is noted from Figure 14 that the maximum allowable recycle ratio is 0.85. Figure 14 also shows that the whole plant would be filled with solids if the recycle ratio were increased to 0.91. This is a theoretical estimate. The operation of the real system would become impracticable due to unmanageable hydrodynamic conditions and accumulation of solids in the reactor for a much lower (than 0.91) value of the recycle ratio. Therefore, the range of values for 0between 0.855 and 0.91 is shown as an impractical range of operation. It is noted that due to the sensitivity of solids concentration to recycle ratio, 8, the maximum liquid production case in Figure 13 almost represents the borderline separating practical and impractical ranges of operation.
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 1, 1983 I
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Figure 15. Variation in liquid yield and liquid production rate as functions of coal concentration in feed, under conditions of high production rate.
H. Process Conditions for High Liquid Production Rate. Based on the studies of the effects of various variables on liquid yield and production rate, a set of conditions is selected to obtain a high liquid production rate. Resulta in Figure 15 show the variations in liquid yield and liquid production rate as a function of coal concentration in feed. The behavior of liquid yield is similar to that observed in the first set of conditions in Figure 12. Production rate of liquid shows a maximum at 26 w t % coal concentration in feed, similar to the one obtained for the first set of conditions at 23.5 wt % coal in feed. The liquid yield in Figure 15 also shows a minimum at feed coal concentration of 36 wt %. Probably for the first set of conditions, also, such a minimum could be observed beyond 40 wt % coal in feed. However, the liquid production rate for the condition in Figure 15 is very high (36.8% larger) compared to that under the first set of conditions. Also, the level of solids in the vacuum column bottoms and hydrogen consumption, wt % mf coal, is lower in this m e (Figure 16) compared to that under the first set of conditions. Comparison of the two sets of conditions presented in Figures 12 and 15 clearly shows that the production rate of liquid can be increased with a simultaneous decrease in hydrogen consumption and solids concentration in vacuum tower bottoms, by increasing reactor temperature and decreasing slurry residence time. The liquid production rate should be expected to be limited by the maximum allowable temperature of the reactor and the preheater capacity. However, with these limitations the liquid production rate can be increased by more than 40 wt % from that normally achieved. In case the objective for the process is maximization of liquid yield, two important limitations of the process must be considered. The first limitation is imposed by the requirement that vacuum tower bottoms be gasified to obtain the required hydrogen for the process. However, it is obvious that the basis for such a dependence on vacuum tower bottoms lies in the cost associated with the separation of SRC from solids. Otherwise, refined coal is more valuable than the raw coal. Therefore, if all of the SRC could be converted to oil, raw coal could be gasified to generate the required amount of hydrogen. However,
25 20
I 25
,
1 I
1
1
30
35
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Coal Concentrstlon i n Feed (o,coal), Wt%
Figure 16. Solids concentration in vacuum column bottoms and hydrogen requirements of high productivity case.
larger amounts of hydrogen consumed in additional conversion of coal may be considered as undesirable if it is not noted that most of it is used in the generation of gases, mainly light hydrocarbon gases. Therefore, if the conversion to liquid were almost complete, a large amount of light hydrocarbon gases would be generated which could be relatively easily converted back to hydrogen by steam reforming. The process would still be self-sufficient in hydrogen; however, a detailed study would be required to evaluate the merits and demerits of such a change in process design. If the conversion of SRC to liquids is nearly complete, separation of.liquids from solids would require a different separation scheme and therefore operability of vacuum column may not be a limitation. Thus, from the above variable studies it is noted that there are innumerable possibilities of increasing liquid yield from the SRC-I1 process by changing the operating conditions within a relatively narrow range. Also, the process offers a large number of alternatives which can be utilized to achieve almost complete conversion of maf coal to liquids. Although the nature of the processes leading to the formation of insoluble organic matter is not clearly understood, it can be speculated that the yield, wt % mf coal, or production of insoluble organic matter would be very small under the high liquid yield conditions. Conclusions Compared to the conditions normally employed in SRC-I1 pilot plants, a 3-6 w t % increase in liquid yield (wt % mf coal) as well as liquid production rate [kg/(L h)] can be achieved by increasing the reactor temperature by 10 "C (455 to 465 "C) or total solids in feed by 5 wt 70 (45 to 50 wt %). By decreasing the coal concentration by 5 wt % (30 to 25 wt %), the liquid yield can be increased by 6-8 wt% mf coal; however, the increase in liquid production rate would be within 1%. Liquid production rate can be increased by about 8% by increasing the slurry feed rate by 30%, i.e., by decreasing the nominal slurry residence time by about 23% (1.108 to 0.852 h). By selecting a proper set of conditions, the liquid production rate could be increased by more than 1.4 times that achieved in the pilot plants. Conditions leading to almost complete conversion of liquefiable organic matter in coal to liquid and
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gaseous products can also be selected. However, in any such case, very high solids contents (greater than 50 wt % ) in the vacuum flash column bottoms would require a different separation scheme, and the amount of hydrogen consumed would be very high (greater than 8 w t % mf coal), whereas little SRC would be produced. Therefore, hydrogen for the process would have to be generated from some other raw material. Nomenclature A's = parameters in gas holdup correlation uo,u = gas-liquid interfacial area in an agitated vessel for low and high levels of mixing, cm-' ASH = inorganic mineral matter cy = oil recycle ratio B = flow rate of bottoms from recycle splitter, kg/h bl,, b, = parameters in vapor pressure correlation p = bottoms (recycle splitter) recycle ratio C = number of components in a flash separator D = flow rate of recycle or distillate oil, kg/h dim= impeller diameter, cm e, = absolute and average errors, percent e, = gas holdup in a bubble column = gas holdup in a stirred vessel - solid holdup F = ASH and IOM free feed to a separator, kmol/h f;"' = fraction of coal going to component i during instantaneous dissolution fFRC = fraction of component i in SRC reaction products G = mass flow rate through the reactor, kg/h G (m,n) = mass flow rate in stream n to or from unit m, kg/h gASH = mass fraction of ASH in the reactor g, (m,n)= mass fraction of component i in stream n from or to unit m,kg/h g, = mass fraction of solids in feed g,, gf,= mass fraction of component i in reactor and feed, respectively H5= vacuum flash bottoms rate, kg/h HD = heavy distillate (288-482 "C) IOM = pyridine insoluble organic matter K,= equilibrium vaporization constant for component i LF = flow rate of flash liquid product, kmol/h L, = flow rate of lighter components from separator n,kg/h LD = light distillate (c5-193 "C) A, = yield of component i, wt % mf coal MD = middle distillate (193-288 "C) mf = moisture-free m, = molecular weight of component i pL = viscosity of liquid, CP vL = kinematic viscosity of liquid N = impeller speed, rpm P, = power dissipated per unit volume, hp/ft3 P = pressure, MPa p H 2= partial pressure of hydrogen, MPa q = total number of data points R, = universal gas constant, kJ/(kmol K) R = rate of recycle, kg/h rSRC = rate of SRC reaction, kg/(L h) p , = density of component i, kg/L pL = density of liquid, kg/L
2-
ps = density of the slurry, kg/L u = liquid surface tension, dyn/cm
SRC = solvent refiied coal (pyridine soluble organic matter boiling above 482 "C) T = temperature, K ,Z' = boiling temperature, "C 7 = nominal slurry residence time, h u, = superficial gas velocity, cm/s ut = bubble rise velocity, cm/s V , = flow rate of flash vapor product, kmol/h VR = volume of the reactor, L VP, = vapor pressure of component i u, = volume fraction of component i x , = mole fraction of component i in feed to a flash column xF, = mole fraction of component i in flash liquid product yFI = mole fraction of component i in flash vapor product Superscripts D = distillate H5= vacuum column bottom product Lz = gas and water separator L4 = fractionator R = recycle slurry Subscripts ASH = inorganic mineral matter B = recycle splitter bottom c = any coal F = flash product f = feed to the reactor G = gas
IOM = pyridine insoluble organic matter i = component i L = liquid phase 0 = standard set of conditions Pow = Powhatan No. 5 coal having 10.1 w t % ASH v = vapor
Literature Cited Akita, K.; Yoshida. F. Ind. Eng. Chem. Process Des. D e v . 1873, 12, 76. Calderbank, P. H. Trans. Inst. Chem. Eng. 1868, 36, 443. Deckwer, W. D. "Access of Hydrodynamic Parameters Required In the Design and Scale-up of Bubble Column Reactws", presented at the Second Chemical Congress of the North American Continent, San Francisco,CA, Aug I 9 8 0 INDE, No. 35. Gopal, J. S.; Shah, Y. T.; Can,N. L. "Estimation of Vapor-Uquid Stream Compodtion in the seperatbn Unit of an SRC-I1 Plant", Report No. 627RM098, Chemicals and Mlnerals Dhriskn, Qulf Research 8 Develop ment Co.,Plttsbwgh, PA, 1981; also DOE-ET-1010448. Kara, S.; Kelkar. 0. G.; Shah, Y. T. Ind. Eng. Chem. Process Des. Dev. 1882. 21, 584. "Process Design Specifications, Coal Diaeohrer Unit (Unit 11) of the 8000 TPSD SRC-I1 Demonstration h n t " , Technobgy and Matsrlels Department, Gulf Research 8 Development Co., Pittsburgh, PA, Report No. 58ORL129. June 1980. ... Singh, C. P. P.:-Sheh, Y . T.; Carr, N. L.; Prudlch, M. E. Can. J . Chem. Eng. 1802a. 60. 248. Singh, C.'P. P.; Shah, Y. T.; Can. N. L. Chem. Eng. J . 1082b, 23, 101. Starling, K. E.; Brute, M. R.; Un, C. T.; Watanaslri, S. "Calculation of Distillable Coal Fluid Themphysicel Properties Using Murtiparameter Conespondkrg States Correlations", Repat No. OU/IQT/S-14366/1, School of Chem. Mat. Science. University of Oklahoma, Norman, OK, Aug 1980. ~
Receiued for review February 24, 1982 Revised manuscript received July 2, 1982 Accepted July 12, 1982
