single-cycle mixed-fluid lng (prico) process
DESCRIPTION
Single-cycle mixed-fluid LNG (PRICO) process. Part I: Optimal design Sigurd Skogestad & Jørgen Bauck Jensen Quatar, January 2009. Single-cycle mixed fluid LNG (PRICO) process. 45 kg/s 30 °C 40 bar. Natural gas : 45 kg/s (1.3 MTPA) Feed at 40 bar and 30 °C - PowerPoint PPT PresentationTRANSCRIPT
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Single-cycle mixed-fluid LNG (PRICO) process Part I: Optimal designSigurd Skogestad & Jørgen Bauck JensenQuatar, January 2009
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Single-cycle mixed fluid LNG (PRICO) process
Natural gas:• 45 kg/s (1.3 MTPA)• Feed at 40 bar and 30 °C
– 89.7 mol% C1, 5.5% C2, 1.8% C3, 0.1% C4, 2.8% N2
• Cooled to ~ -156 °C• Expansion to ~ 1 bar
– Flash gas may be used as fuel
• Liquefied natural gas (LNG) product at -162C
-162 °C
1 bar
45 kg/s30 °C40 bar
-156 °C 35 bar
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Single-cycle mixed fluid LNG (PRICO) process
Refrigerant:• Mixed fluid: ~ 33mol% C1, 35%
C2, 25% C4, 7% N2
• Partly condensed with sea water to ~ 30 °C
• Cooled to ~ -156 °C• Expansion to ~ 4 bar• Evaporates in NG HX• Super-heated ~ 10 °C• Compressed to ~ 22 bar
4 barSup 10 °C
-156°C 19 bar
22 bar45 kg/s30 °C40 bar
475 kg/s30 °C22 bar
-156 °C
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Compressor:• Max. pressure: 22 bar / 30 bar• Max. compressor suction
volume*: 317000 m3/h• Max. compressor head*: 263.6
kJ/kgOr: Max. compressor ratio* Pr, e.g. 5.5 (Price)
4. Max. compressor work: 77.5 MW / 120 MW
5. Minimum superheating: 10C
Design constraints
-162 °C
30 °C40 bar
1 bar
3.33 kg/s * Design constraint only
30 °C
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Optimal design: TAC
min JTAC = Joperation + Jcapital
subject to c ≤ 0• Joperation [$/year] is the annual operating cost
– Joperation = Jutility + Jfeeds + Jproducts
• Jcapital [$/year] is the annualized cost of the equipment
• Total annualized cost (TAC) is minimized with respect to the design variables
– Flowsheet structure – Areas, sizes– Operating parameters (pressures etc.)
• Requires mixed integer non-linear programming• Our case Fixed structure Try a simpler approach
Maximize total profit = Minimize Total Annualized Cost (TAC):
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Idea: Specify ΔTmin to balance between
• operating costs (favoured by a low value)
• capital costs (favoured by a high value)
Simpler approach: Specify ΔTmin
-162 °C
30 °C40 bar
1 bar
3.33 kg/s * Design constraint only
ΔTmin=2C*
30 °C
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Simple ΔTmin-method (Approach 1)
• ΔTmin (=2C) is added as an extra design constraint + minimize compressor work (Ws)
• BUT: The resulting design parameters (pressure etc.) are not optimal for the resulting process! – Reoptimizing reduces ΔTmin to about 1C and reduces work by about
5% (!)– Cannot be fixed by iterating on ΔTmin
• Therefore: Approach 1 NOT USED
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Simplified TAC (sTAC)
Capital costJcapital = Σi (Cfixed,i + Cvariable,i·Si
ni) / TT – capital depriciation time, e.g. 10 years
1. Structure of plant given Cfixed,i = 0
2. Main equipment: Heat exchangers and compressor 3. Scaling exponent
• n = 1 for compressor (Can then combine operation and capital cost!)• n = 0.65 for heat exchangers
4. Cvariable,i = C0 for all heat exchangers
Approach 2: Adjust C0 to get ΔTmin = 2C
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sTAC – Optimization problemMinimize cost
Case I: Feedrate (NG) givenCase II: Feedrate free
Here: Consider Case II.Minimize cost=”Max. single-train LNG feed”
3.33 kg/s
1 bar
30 °C40 bar
30 °C
-162 °C
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Resulting “Max feed” sTAC:
• Minimization with respect to design parameters (AHOT
and ANG) and operating parameters (pressures etc.)– ANG: NG / cold refrigerant
– AHOT: hot refrigerant / cold refrigerant
• Here: Adjust C0 to obtain ΔTmin = 2C• Other constraints c: depend on specific case
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Case 1 – Price and Mortko (1983)• Data
– LNG outlet temperature (before expansion) = -144 °C– 77.5 MW compressor power– Maximum Ph = 22 bar
– Maximum Pr = Ph/Pl = 5.5
• Differences / uncertainties– Pure methane – Neglected removal of heavy components– Pressure losses (especially important at low pressure, e.g. compressor
suction)– Heating of fuel gas produces some LNG “for free”
• 3.7 % higher production compared with Price & Mortko– 44.6 kg/s LNG production– Gives large amount of fuel gas (7.7 kg/s, ~230 MW)
• Want to limit fuel to 3.33 kg/s, ~100 MW
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Case 2 – Limited fuel flow• Limitation on fuel flow instead of outlet temperature
– Maximum 3.33 kg/s of fuel (7.7. kg/s in Case 1)– Outlet temperature down from -144 °C to -156 °C to get sufficient
cooling with less flash gas (fuel)– Production (with Ws=77.5 MW and Pr=5.5) reduced by 6 % compared
with case 1• From 44.6 kg/s to 41.7 kg/s
3.33 kg/s
77.5 MW
-162C41.7 kg/s
-156C
22 bar
4 bar
45 kg/s30C
475 kg/s30C
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Case 3,4 – Super-heating• Wish to find the optimal degree of super-heating
– 10.0 °C super-heating used for all cases except 3 and 4– Case 3; 11.6 °C super-heating increases production by 0.8 % compared with
case 2– Case 4; 25.7 °C super-heating decreases production by 1.3 % compared with
case 3
• Optimum is very flat in terms of super-heating• Some super-heating is necessary to protect the compressor• Some super-heating is optimal due to
– Internal heat exchange in the main heat exchanger
• However, the heat transfer coefficient in the super-heating region is lower than in the evaporating region
– This has not been considered here– Will tend to reduce the optimal amount of super-heating
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Case 5 – No pressure constraint• We have removed the following constraints
– Maximum Ph = 22 bar– Maximum Pr = Ph/Pl = 5.5
• Ph is increased to 50.4 bar and Pr is increased to 22• LNG production is increased by 11 % (from case 2)• The high pressure ratio is not possible with a single compressor
casing– The compressor head is too high– Two compressors in series will do the job
• Higher head [kJ/kg] gives lower refrigerant flow– Cooling duty per kg of refrigerant closely related to head– Less heat transfer area is needed since less warm refrigerant needs cooling
• The cost of an additional compressor casing is at least partly offset by the decreased heat transfer area and increased production
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Case 6,7 – Real GE Compressor• GE MCL1800 series compressor
– Centrifugal compressor with 1800 mm casing diameter– Maximum suction volume is 380 000 m3/h active constraint– Maximum discharge pressure Ph = 30 bar active constraint
• Case 6 – 77.5 MW; Same production as case 5 – Compressor head is 216 kJ/kg which is feasible with a single
compressor casing
• Case 7 – 120 MW; 71.1 kg/s of LNG product– Compressor head is 162 kJ/kg which is feasible with a single
compressor casing– Corresponds to 2.0 million tons per annum (MTPA) with 330
operating days per year
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Case 8 – Liquid turbines• Expansion in liquid turbines
– Takes the pressure down to 2 bar above the saturation pressure
– Avoid vapour in the turbines– Possible with two phase turbines?
• Production increased by 6.6 % compared with case 7– 75.8 kg/s ~ 2.2 MTPA per train
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Production vs. feed pressure• Results for case 8• Achievable feed
pressure depends on – Location of heavy extraction
• Up-front or integrated• Recompression after heavy
extraction– Feed compressor?
• Complicates the optimization problem– Very important for production
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Conclusion
• The constraints on the compressor performance is very important for the maximum production design case
– Maximum compressor head– Maximum compressor shaft work– Maximum compressor suction volume
• The feed pressure is very important for the achievable production
– We have assumed a fixed feed pressure of 40 bar
• A large PRICO train of 2.2 MTPA is feasible with a single compressor casing
– 2.0 MTPA without liquid turbines
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Conclusion II
• All the results presented here are with a minimum approach temperature ΔTmin = 2.0 °C– This is achieved by adjusting C0 in the optimization problem
• An alternative is to find a reasonable C0 and the use the same value for all cases– These results are presented in the paper
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Additional material
1. Table with results for all cases2. Table with results for the alternative design method
with constant C0
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Fixed C0 for all cases