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Vapor Pressure of crude oil
- Thread starter asade
- Start date
RVP is a theretical value for comparrion of different crudes. Cant use that for sizing. TVP is temperature specific, assuming that you know it at your operating pressure this is what you need to supply to the vendor.
Best regards
Morten
PS: I am what i am through a combination effort and luck
Best regards
Morten
PS: I am what i am through a combination effort and luck
ooops a couple of typoes that i feel that i need to correct
RVP is a theoretical value for comparrion of different crudes. Cant use that for sizing. TVP is temperature specific, assuming that you know it at your operating _temperature_ this is what you need to supply to the vendor.
Best regards
Morten
RVP is a theoretical value for comparrion of different crudes. Cant use that for sizing. TVP is temperature specific, assuming that you know it at your operating _temperature_ this is what you need to supply to the vendor.
Best regards
Morten
- Thread starter
- #4
If you dont already know it there are a number of ways:
1) Measure it - for a stable supply of crude this is the most reliable methode. The info should come with a crude assay that is commonly supplied at the delivery to the refinery
2) Estimate it by thermodynamics (e.g. a simulation tool such as HYSYS). The reliability of this will of course depend on your input data
3) API and other may have standard curves with a few input parameters e.g. API gravity (std. density). I dont know exactly where to look but try to search for API databook or similar. If you are working for an oil company they may also have similar data.
4) It could perhaps be derived from upstream conditions? E.g. if the valve sits on the outlet from a separator then the TVP=operating pressure of the vessel
Best regards Morten
1) Measure it - for a stable supply of crude this is the most reliable methode. The info should come with a crude assay that is commonly supplied at the delivery to the refinery
2) Estimate it by thermodynamics (e.g. a simulation tool such as HYSYS). The reliability of this will of course depend on your input data
3) API and other may have standard curves with a few input parameters e.g. API gravity (std. density). I dont know exactly where to look but try to search for API databook or similar. If you are working for an oil company they may also have similar data.
4) It could perhaps be derived from upstream conditions? E.g. if the valve sits on the outlet from a separator then the TVP=operating pressure of the vessel
Best regards Morten
What are the parameters for the control valve? Depending what the inlet and outlet pressures are the the vapor pressure may not be critical and in many cases, you should be able to come up with a reasonable guesstimate by looking at the process.
Fisher's control book has a good section on cavitation and how to use it to correct for it with the standard liquid sizing formula. In lots of cases, the cavitation doesn't really impact the valve sizing but is an indication you may want to look at hardened trim to prevent internal damage.
For example, if you were sizing a minimum flow control valve for a crude charge pump, you might be sizing the valve with 300 psig or more on the inlet and a few psig on the outlet if you are going back to a storage tank. In that case, I'd spend more time coming up with a reasonable vapor pressure than if I was sizing a control valve for the crude train where I have 300 psig inlet and 250 psig outlet. In the latter case, the vapor pressure isn't a significant factor.
I'd be cautious using Hysys. The components that drive the vapor pressure are the light components that are the smallest percentages and the easiest to lose during sample collection and analysis. I've seen liquid samples that would be at their bubble point taken from vessels that when you put the composition back into Hysys as a bubble point liquid with either a pressure or temperature, Hysys doesn't come close to matching the temperature or pressure. Similar comments go for vapor samples but where you are trying to calculate a dewpoint.
Fisher's control book has a good section on cavitation and how to use it to correct for it with the standard liquid sizing formula. In lots of cases, the cavitation doesn't really impact the valve sizing but is an indication you may want to look at hardened trim to prevent internal damage.
For example, if you were sizing a minimum flow control valve for a crude charge pump, you might be sizing the valve with 300 psig or more on the inlet and a few psig on the outlet if you are going back to a storage tank. In that case, I'd spend more time coming up with a reasonable vapor pressure than if I was sizing a control valve for the crude train where I have 300 psig inlet and 250 psig outlet. In the latter case, the vapor pressure isn't a significant factor.
I'd be cautious using Hysys. The components that drive the vapor pressure are the light components that are the smallest percentages and the easiest to lose during sample collection and analysis. I've seen liquid samples that would be at their bubble point taken from vessels that when you put the composition back into Hysys as a bubble point liquid with either a pressure or temperature, Hysys doesn't come close to matching the temperature or pressure. Similar comments go for vapor samples but where you are trying to calculate a dewpoint.
- Thread starter
- #7
TD2K,
One of the reason I am worried about the vapour pressure is large pressure drop across the valve. For instance, I have a LV on a recirculation line of a crude oil pump. The pump discharges at 526 psig. The recirculation line was connected to a 1.50 psig surge tank. The pressure drop across the valve is expected to be 524.5 psi.
HYSYS only provides TVP and RVP properties of the crude stream. My challenge is - can TVP be assumed as the vapour pressure for the crude oil for the sake of the LV? I researched that the fluid pressure should not fall below the vapour pressure of the fluid to prevent bubble formation in the liquid. HYSYS parameters is the ONLY information available to me now.
How do I go about it?
I am what I am by His grace
One of the reason I am worried about the vapour pressure is large pressure drop across the valve. For instance, I have a LV on a recirculation line of a crude oil pump. The pump discharges at 526 psig. The recirculation line was connected to a 1.50 psig surge tank. The pressure drop across the valve is expected to be 524.5 psi.
HYSYS only provides TVP and RVP properties of the crude stream. My challenge is - can TVP be assumed as the vapour pressure for the crude oil for the sake of the LV? I researched that the fluid pressure should not fall below the vapour pressure of the fluid to prevent bubble formation in the liquid. HYSYS parameters is the ONLY information available to me now.
How do I go about it?
I am what I am by His grace
- Thread starter
- #8
The 526 psig above is the operating discharge pressure from the pump. The recirculation flow is expected to occur at a minimum flow of the pump which could occur at a pressure of 630 psig due to low throughput.
From HYSYS, the RVP and TVP at 38 deg.C is obtained as -10.55 psig and 0.1174 psig respectively. This condition was obtained at the pump discharge outlet on the HYSYS model.
I am what I am by His grace
From HYSYS, the RVP and TVP at 38 deg.C is obtained as -10.55 psig and 0.1174 psig respectively. This condition was obtained at the pump discharge outlet on the HYSYS model.
I am what I am by His grace
- Thread starter
- #9
Crudes do not have high vapor pressures. What they do have is mostly attributable to a little entrained gas. Gasoline VP is normally less than 9 psia.
"People will work for you with blood and sweat and tears if they work for what they believe in......" - Simon Sinek
"People will work for you with blood and sweat and tears if they work for what they believe in......" - Simon Sinek
asade:
As i wrote: If you last stage separator is a "flash" (that is upstream has higher pressure) - then BY DEFINITION your crude will leave the vessel at its boiling point. So in your case: 55 psig at 38ºC. I must also say that this is a high operating pressure for a last stage separator. Cool the fluid and the VP will drop.
Best regards
Morten
As i wrote: If you last stage separator is a "flash" (that is upstream has higher pressure) - then BY DEFINITION your crude will leave the vessel at its boiling point. So in your case: 55 psig at 38ºC. I must also say that this is a high operating pressure for a last stage separator. Cool the fluid and the VP will drop.
Best regards
Morten
As Morten said, the 55 psig is the crude vapor pressure, I'm not sure how you have your simulation set up if Hysys is saying the vapor pressure of the crude is 0.1174 psig. Given your description of the process, you would have to further process that oil to get it down to a TVP of 0.1174 psig.
If you route that oil to an atmospheric tank (I'm assuming the 1.5 psig is a static head in the tank) you will have vapors flashing off the crude oil. You'll have flashing through your control valve and possibly a safety/environmental problem with the vapors that will come off that surge tank's vents.
Why don't you recirculate on low flow back to the separator operating at 55 psig?
If you route that oil to an atmospheric tank (I'm assuming the 1.5 psig is a static head in the tank) you will have vapors flashing off the crude oil. You'll have flashing through your control valve and possibly a safety/environmental problem with the vapors that will come off that surge tank's vents.
Why don't you recirculate on low flow back to the separator operating at 55 psig?
- Thread starter
- #14
TD2K,
For clarity, the RVP and TVP at 38 deg.C from the crude storage tank in simulation was obtained as -10.55 psig and 0.1174 psig respectively, which is also the values obtained at the pump discharge line. Would it be save to connect the pump recycle to the surge tank (operating pressure @ 1.50 psig).
The 1st stage separator is operating at 60 psig (not 55 psig) and its crude oil outlet TVP at 38 deg.C was 55.07 psig.
Now, can I safely take the TVP values as the vapour pressure for the crude oil?
I am what I am by His grace
For clarity, the RVP and TVP at 38 deg.C from the crude storage tank in simulation was obtained as -10.55 psig and 0.1174 psig respectively, which is also the values obtained at the pump discharge line. Would it be save to connect the pump recycle to the surge tank (operating pressure @ 1.50 psig).
The 1st stage separator is operating at 60 psig (not 55 psig) and its crude oil outlet TVP at 38 deg.C was 55.07 psig.
Now, can I safely take the TVP values as the vapour pressure for the crude oil?
I am what I am by His grace
Asade, give us a sketch of your system please. I'm not quite sure what we are talking about and I don't want to lead you astray. It would help if you can note on the sketch, if it's not clear, where the various above data/values apply.
I'd be leary of a vapor pressure from Hysys if it's from a lab analysis for the above reasons. The vapor pressure of a crude oil stored in a floating roof tank is likely going to vary from a few psia up to about 11 psia. If you have a flash point for the crude you may be able to estimate it off that. Heavier hydrocarbons have a LEL of about 1 to 1.5 mol%. So, your flash point essentially gives you a vapor pressure of about 0.15 psia at the flash temperature which you can then correct to another temperature using some of the curves in the API data book. It's not great but it will give you another check against what Hysys gives you.
Run your valve sizing program with a range of vapor pressures (TVP) and see if you have a problem.
I'd be leary of a vapor pressure from Hysys if it's from a lab analysis for the above reasons. The vapor pressure of a crude oil stored in a floating roof tank is likely going to vary from a few psia up to about 11 psia. If you have a flash point for the crude you may be able to estimate it off that. Heavier hydrocarbons have a LEL of about 1 to 1.5 mol%. So, your flash point essentially gives you a vapor pressure of about 0.15 psia at the flash temperature which you can then correct to another temperature using some of the curves in the API data book. It's not great but it will give you another check against what Hysys gives you.
Run your valve sizing program with a range of vapor pressures (TVP) and see if you have a problem.
- Thread starter
- #16
TD2K,
The thread started because I was asked to provide the vapour pressure of the fluid that passes through the LV valves. Majority of these LVs control crude oil flows.
Example is in the attached drawing that shows an LV-8430 on the recirculation line from the pipeline pumps. The vapour pressure of the crude oil flowing through the valve shoud be provided. From the pump discharge line, RVP and TVP were obtained as 10.55 psig and 0.1174 psig respectively from the HYSYS model. Can TVP be used?
I hope I have been able to clear some doubts.
I am what I am by His grace
The thread started because I was asked to provide the vapour pressure of the fluid that passes through the LV valves. Majority of these LVs control crude oil flows.
Example is in the attached drawing that shows an LV-8430 on the recirculation line from the pipeline pumps. The vapour pressure of the crude oil flowing through the valve shoud be provided. From the pump discharge line, RVP and TVP were obtained as 10.55 psig and 0.1174 psig respectively from the HYSYS model. Can TVP be used?
I hope I have been able to clear some doubts.
I am what I am by His grace
It seems to me that the pressure in the tank that the pumps draw from must be atmospheric.
SO - again by definition the VP will be around atmospheric at whatever storage temperature you have - or lower depending on upstream conditions. If you process model says its 55 psig@38ºC at the inlet to this tank then you must experience a lot of flashing of (degassing) from this tank. Its not obvious since the ventheader seems to be moderate in size (i read 6" or 8") whereas the pumps are pretty big at 2000 GPM. This fits with what you write at your TVP now (0.1174) psig - dont really see how the 55 pisg originally came into the picture? So downstream the pumps the oil will be a bit hotter trhan in the tanks and the TVP thus a bit higher than atmosperic. So yes, i would go with the calculated TVP for valve specification any time of the day
Best regards
Morten
SO - again by definition the VP will be around atmospheric at whatever storage temperature you have - or lower depending on upstream conditions. If you process model says its 55 psig@38ºC at the inlet to this tank then you must experience a lot of flashing of (degassing) from this tank. Its not obvious since the ventheader seems to be moderate in size (i read 6" or 8") whereas the pumps are pretty big at 2000 GPM. This fits with what you write at your TVP now (0.1174) psig - dont really see how the 55 pisg originally came into the picture? So downstream the pumps the oil will be a bit hotter trhan in the tanks and the TVP thus a bit higher than atmosperic. So yes, i would go with the calculated TVP for valve specification any time of the day
Best regards
Morten
- Thread starter
- #18
the vapor pressure for a mixture is the bubble point pressure at specified temperature
( see ) ,
there are several methods to determine the True Vapor Pressure, which depends from composition and temperature, see for example ASTM.
Keep in account that many manufacturers utilize the ISA formulation for control valves, in ISA formulation the vapor pressure is that of pure fluids, for mixtures you may provide an average value or some equivalent parameter.
( see ) ,
there are several methods to determine the True Vapor Pressure, which depends from composition and temperature, see for example ASTM.
Keep in account that many manufacturers utilize the ISA formulation for control valves, in ISA formulation the vapor pressure is that of pure fluids, for mixtures you may provide an average value or some equivalent parameter.
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